Methods and systems for upgrading heavy oil using catalytic hydrocracking and thermal coking

ABSTRACT

Methods and systems for hydroprocessing heavy oil feedstocks to form upgraded material use a colloidal or molecular catalyst dispersed within heavy oil feedstock, pre-coking hydrocracking reactor, separator, and coking reactor. The colloidal or molecular catalyst promotes upgrading reactions that reduce the quantity of asphaltenes or other coke forming precursors in the feedstock, increase hydrogen to carbon ratio in the upgraded material, and decrease boiling points of hydrocarbons in the upgraded material. The methods and systems can be used to upgrade vacuum tower bottoms and other low grade heavy oil feedstocks. The result is one or more of increased conversion level and yield, improved quality of upgraded hydrocarbons, reduced coke formation, reduced equipment fouling, processing of a wider range of lower quality feedstocks, and more efficient use of supported catalyst if used with the colloidal or molecular catalyst, as compared to a conventional hydrocracking process or a conventional thermal coking process.

BACKGROUND OF THE INVENTION

1. The Field of Invention

The present invention is in the field of upgrading hydrocarbonfeedstocks that include a significant quantity of asphaltenes andhydrocarbons having a boiling point above 524° C. (975° F.) into lowerboiling, higher quality materials.

2. The Relevant Technology

World demand for refined fossil fuels is ever-increasing and willinevitably outstrip the supply of high quality crude oil, whether as aresult of actual shortages or due to the actions of oil cartels. Ineither case, as the price or shortage of crude oil increases there willbe increasing demand for ways to better exploit lower quality feedstocksand extract fuel values therefrom. As more economical ways to processlower quality feedstocks become available, such feedstocks may possiblycatch, or even surpass, higher quality crude oils as the primary sourceof refined fossil fuels for operating automobiles, trucks, farmequipment, aircraft, and other vehicles that rely on internalcombustion.

Low quality feedstocks include relatively high quantities ofhydrocarbons that have a boiling point of 524° C. (975° F.) or higher.They also contain relatively high concentrations of sulfur, nitrogen andmetals. High boiling fractions typically have a high molecular weightand/or low hydrogen/carbon ratio and include complex compoundscollectively referred to as “asphaltenes”. Asphaltenes are difficult toprocess and commonly cause fouling of conventional catalysts andhydroprocessing equipment.

Examples of lower quality feedstocks that contain relatively highconcentrations of asphaltenes, sulfur, nitrogen and metals include heavyoil and oil sands bitumen, as well as bottom of the barrel and residuumleft over from conventional refinery process (collectively “heavy oil”).The terms “bottom of the barrel” and “residuum” (or “resid”) typicallyrefer to atmospheric tower bottoms, which have a boiling point of atleast 343° C. (650° F.), or vacuum tower bottoms, which have a boilingpoint of at least 524° C. (975° F.). The terms “resid pitch”, “vacuumresidue” and “vacuum reduced crude” (VRC) are commonly used to refer tofractions that have a boiling point of at least 524° C. (975° F.).

By way of comparison, Alberta light crude contains about 9% by volumevacuum residue, while Lloydminster heavy oil contains about 41% byvolume vacuum residue, Cold Lake bitumen contains about 50% by volumevacuum residue, and Athabasca bitumen contains about 51% by volumevacuum residue. Resid contains even higher concentrations of fractionsthat boil at or above about 343° C. (650° F.), with vacuum tower bottomsalmost exclusively comprising fractions that boil at or above about 524°C. (975° F.).

In conventional petroleum refining processes, crude oil typically isfractionated by an atmospheric distillation tower, producing fractionswith different boiling points, including: gases, light naphtha, heavynaphtha, jet fuel, kerosene, diesel oil, atmospheric gas oil, andatmospheric bottoms (or atmospheric reduced crude). Among theseproducts, gases undergo gas processing that eventually yields productsincluding fuel, butanes, liquefied petroleum gas (LPG), and the like.The most commercially valuable fractions are the lower boiling liquidfractions, which undergo further hydroprocessing, includinghydrocracking and hydrotreating, to yield gasoline blending products,jet fuel, kerosene, and diesel oil. The highest boiling fractions,atmospheric bottoms, are further fractionated by a vacuum distillationtower, producing fractions with increasing boiling points including:gas, light vacuum gas oil, heavy vacuum gas oil, vacuum residuum (orvacuum reduced crude), and asphalt. Light vacuum gas oil and heavyvacuum gas oil are further processed to yield gasoline blendingproducts, while vacuum residuum is typically further processed by acoker, i.e., a system that reforms high boiling heavy oil (typicallyvacuum residuum) by thermal cracking, forming upgraded hydrocarbons andcoke.

Coking is a thermal cracking process used in oil refineries to upgradeand convert petroleum residuum (bottoms from atmospheric and vacuumdistillation of crude oil) into liquid and gas product streams, leavingbehind a solid concentrated carbon material, petroleum coke. Cokingproducts include gas, coker naphtha, coker gas oil, and petroleum coke,among which coker naphtha and coker gas oil are the more commerciallyvaluable fractions, and can be further processed to yield kerosene,diesel, and gasoline blending products. According to the structure ofpetroleum coke, coke products include needle coke, sponge coke, shotcoke, and anode grade coke.

Converting heavy oil into useful end products requires extensiveprocessing, including reducing the boiling point of the heavy oil,increasing the hydrogen-to-carbon ratio, and removing impurities such asmetals, sulfur, nitrogen and high carbon forming compounds. Examples ofcatalytic hydrocracking processes using conventional supported catalyststo upgrade atmospheric tower bottoms include slurry bed hydroprocessingthat utilizes fine solid catalyst particles, fixed-bed hydroprocessingthat utilizes a solid heterogeneous catalyst, ebullated- or expanded-bedhydroproces sing that utilizes a solid heterogeneous catalyst, andmoving-bed hydroproces sing that utilizes a solid heterogeneous catalystand is a version of fixed bed hydroproces sing. Non-catalytic processesused to upgrade vacuum tower bottoms include thermal cracking, such asresidual oil coking (for which several commercial processes existincluding delayed coking, fluid coking and ExxonMobil's proprietaryFlexicoking® process), and solvent extraction. Solvent extraction isquite expensive and incapable of reducing the boiling point of the heavyoil. Conventional catalytic hydrocracking processes often involve rapidcatalyst sintering, fouling, and deactivation and high catalyst cost,making them currently unsuitable for hydroprocessing vacuum towerbottoms unless substantially diluted with lower boiling fractions, suchas atmospheric tower bottoms. Even with dilution of the lower boilingfractions, most existing ebullated bed processes operate at less than 65wt % conversion, while most fixed bed processes have less than about 25wt % conversion. Coking currently is the primary commercial method toupgrade vacuum reduced crude, but conventional coking processestypically are associated with low conversion rate, high risks of cokingand fouling of equipment, extensive decoking time, and high decokingwater and energy consumption.

When coke is formed in hydroprocessing processes other than coking, ittends to foul equipment and deactivate the reactor and catalyst,requiring extensive maintenance, expensive repairs and increasedcatalyst. Even when coke is formed in a coking process, it requiresdecoking using high pressure water and steam in a coking reactor, whichconsumes substantial time, energy, space, and water. Coking also tendsto cause clogging and fouling within the coking system other than thecoking reactor (e.g. furnace, resid feed line, filter, cokingfractionator, and feed line).

Exacerbating the relatively low conversion levels using existinghydroprocessing systems is the inability to proportionally convert theasphaltene fraction at the same conversion level as the heavy oil as awhole. The result of disproportional conversion is progressive buildupof asphaltenes in the processed feedstock, with an attendant increase inthe likelihood that coke and sediment will form in the reactor and otherprocessing equipment. Apart from equipment fouling, coke and sedimentscan lead to instability of residual resid when used as a fuel oil.

In view of the foregoing, there is an ongoing but unsatisfied need todevelop improved hydroproces sing methods and systems that can be usedat a commercial level to upgrade heavy oil feedstocks. There is also anunsatisfied need to develop hydroprocessing methods and systems that canbe used to upgrade vacuum tower bottoms and other low grade heavy oilfeedstocks, increasing the production of upgraded liquid hydrocarbonproducts while reducing the formation of coke.

SUMMARY OF THE INVENTION

The present invention relates to hydroprocessing methods and systems forupgrading heavy oil by combining a hydrocracking process that utilizes acolloidal or molecular catalyst with a coking process. When usedtogether, the two processes increase overall conversion of high boilingfractions into lower boiling fractions and increase C4+ distillateyields.

The methods and systems involve the use of a colloidal or molecularcatalyst dispersed within a heavy oil feedstock, a pre-cokinghydrocracking reactor, and a coking reactor. The colloidal or molecularcatalyst is preferentially associated with asphaltenes and other cokeforming precursors within the heavy oil feedstock, promoting upgradingreactions that reduce the quantity of asphaltenes or other coke formingprecursors, increase the hydrogen to carbon ratio in the upgradedmaterial, decrease the boiling points of hydrocarbons in the upgradedmaterial, increase yield and conversion rate, improve the quality of theupgraded liquid hydrocarbon products, and reduce formation of coke. Themethods and systems can be used to upgrade vacuum tower bottoms andother low grade heavy oil feedstocks.

Conventional coking is the primary commercial method to upgrade vacuumreduced crude because most conventional hydrocracking processes areunsuitable for processing asphaltenes and other coke forming precursors.Instead of directly introducing vacuum reduced crude into a coker, thepresent invention first upgrades the vacuum reduced crude or other heavyoil feedstock in a pre-coking reactor with colloidal-sized particles ormolecules of a hydroproces sing catalyst dispersed throughout thefeedstock, a significant portion of the catalyst being associated withasphaltene molecules present in the feedstock. As the asphaltenemolecules form free radicals under hydrocracking temperature, theclosely associated colloidal or molecular catalyst catalyzes a reactionbetween the asphaltene radicals and hydrogen, thereby preferentiallypromoting beneficial upgrading reactions to form smaller hydrocarbonmolecules instead of forming coke and sediment. As a result, theasphaltene fraction found in heavy oil feedstocks can be upgraded intomore usable materials along with the other hydrocarbons in the feedstockrather than simply being a coke and sediment precursor that is, at best,a low-value product in the coking process, or, at worst, a detrimentalbyproduct that can quickly deactivate the catalyst and/or foul theprocessing equipment, requiring substantially greater quantities ofcatalyst and/or costly shut downs and clean-up operations.

The present invention can increase the formation of upgraded liquidhydrocarbons in the coker and reduce formation of coke. This advantagealso reduces the frequency of online-offline coking-decoking cycles ofcoker drums in delayed coking, which require repeated shut down,decoking with high pressure steam and water, and high temperature andpressure cyclings. The lower decoking frequency greatly increasesequipment operating life and reduces operation cost.

One aspect of the present invention involves a method for hydroprocessing a heavy oil feedstock to reduce coke formation and increaseproduction of upgraded liquid hydrocarbon products, comprising: (1)preparing a heavy oil feedstock comprised of a substantial quantity ofhydrocarbons having a boiling point greater than about 343° C.,including asphaltenes or other coke forming precursors, and a colloidalor molecular catalyst dispersed throughout the heavy oil feedstock; (2)introducing hydrogen and the heavy oil feedstock with the colloidal ormolecular catalyst into a pre-coking hydrocracking reactor; (3) heatingor maintaining the heavy oil feedstock at a hydrocracking temperature toform hydrocarbon free radicals from the heavy oil feedstock, thecolloidal or molecular catalyst catalyzing upgrading reactions betweenhydrogen and the hydrocarbon free radicals to yield an upgradedmaterial, the upgrading reactions reducing the quantity of asphaltenesor other coke forming precursors, increasing the hydrogen to carbonratio in the upgraded material, and decreasing the boiling points ofhydrocarbons in the upgraded material compared to the heavy oilfeedstock; (4) transferring the upgraded material, together withresidual colloidal or molecular catalyst and hydrogen, to a separator toseparate gaseous and volatile fractions from a liquid hydrocarbonfraction, the residual colloidal or molecular catalyst being dispersedin the liquid hydrocarbon fraction; (5) introducing at least a portionof the liquid hydrocarbon fraction into one or more coking reactors,causing thermal-cracking of the liquid hydrocarbon fraction to form cokeand upgraded hydrocarbon products; and (6) separating the coke from theupgraded hydrocarbon products.

Another aspect of the invention involves a hydroprocessing system forhydroprocessing a heavy oil feedstock to form coke and upgradedhydrocarbon products, comprising: (1) a heavy oil feedstock comprised ofa substantial quantity of hydrocarbons having a boiling point greaterthan about 343° C. and a colloidal or molecular catalyst dispersedthroughout the feedstock; (2) a pre-coking hydrocracking reactor thatheats or maintains the heavy oil feedstock at a hydrocrackingtemperature together with hydrogen during use in order to convert atleast a portion of higher boiling hydrocarbons in the heavy oilfeedstock to lower boiling hydrocarbons and thereby form an upgradedmaterial, the pre-coking hydrocracking reactor comprised of (i) an inletport at a bottom of the reactor into which the heavy oil feedstock andhydrogen are introduced and (ii) an outlet port at a top of the reactorfrom which the upgraded material, colloidal or molecular catalyst, andhydrogen are withdrawn; (3) a separator that separates gaseous andvolatile fractions from a higher boiling liquid hydrocarbon fraction inthe upgraded material, the separator comprised of (i) an inlet throughwhich the upgraded material is introduced into the separator, (ii) afirst outlet through which the gaseous and volatile fractions arewithdrawn, and (iii) a second outlet through which the liquidhydrocarbon fraction is withdrawn; and (4) one or more coking reactorsconfigured to receive and process the liquid hydrocarbon fraction, theone or more coking reactors being thermal coking reactors, such asresidual oil coking. Examples of commercially available coking processeswhich are suitable for this invention are delayed coking, fluid coking,and the ExxonMobil Flexicoking® processes.

A guard bed may optionally be used in the foregoing methods and systemsto remove metals from the liquid hydrocarbon fraction produced by thepre-coking reactor before introducing the liquid hydrocarbon fractioninto the coking reactor in order to improve the quality of the resultingpetroleum coke. The guard bed contains a solid supported catalyst forhydroproces sing the liquid hydrocarbon fraction, wherein the solidsupported catalyst removes at least a portion of the residual colloidalor molecular catalyst and metal contaminants from the liquid hydrocarbonfraction.

The colloidal or molecular catalyst used within the method and systemsaccording to the invention is typically formed in situ within the heavyoil feedstock prior to, or upon commencing, hydroprocessing of thefeedstock. According to one embodiment, an oil soluble catalystprecursor comprising an organo-metallic compound or complex is blendedwith the heavy oil feedstock containing sulfur bearing molecules andthoroughly mixed in order to achieve a very high dispersion of theprecursor within the feedstock prior to formation of the catalyst.Catalyst precursors may include catalytic metals such as Mo, Ni, Co, W,Fe, V and combinations thereof. Ligands for the metals can include2-ethyl hexanoate, naphthanate, octoate, hexacarbonyl, pentacarbonyl,3-cyclopentylpropionate, cyclohexanebutyric acid, biphenyl-2-carboxylicacid, 4-heptylbenzoic acid, 5-phenylvaleric acid, geranic acid,10-undecenoic acid, dodecanoic acid, octanoic acid, 2-ethylhexanoicacid, naphthanic acid, pentacarbonyl, hexacarbonyl, and the like. Anexemplary catalyst precursor is a molybdenum 2-ethylhexanoate complexcontaining approximately 15% by weight molybdenum. Another exemplarycatalyst precursor is a molybdenum 3-cyclopentylpropionate complexcontaining approximately 15% by weight molybdenum.

In order to ensure thorough mixing of the catalyst precursor within thefeedstock, the catalyst precursor is preferably pre-blended with ahydrocarbon oil diluent (e.g., vacuum gas oil, decant oil, cycle oil, orlight gas oil) to create a diluted precursor mixture, which isthereafter blended with the heavy oil feedstock. The decompositiontemperature of the catalyst precursor is selected so as to besufficiently high so that the catalyst precursor resists prematuredecomposition before intimate mixing of the catalyst precursor withinthe feedstock. Subsequent heating of the feedstock to a temperaturesufficient to decompose the catalyst precursor and cause the release ofhydrogen sulfide from sulfur-bearing hydrocarbon molecules, eitherbefore or upon commencing hydroprocessing, causes the catalyst precursorthat has been intimately mixed with the feedstock to yield individualmetal sulfide catalyst molecules and/or extremely small particles thatare colloidal in size (i.e., less than about 500 nm, preferably lessthan about 100 nm, more preferably less than about 10 nm, even morepreferably less than about 5 nm, and most preferably less than about 1nm).

The benefits resulting from the foregoing methods and systems includeincreased conversion level and upgraded hydrocarbon yield and quality,reduced coke formation, lowered decoking frequency, decreased energy andwater consumption, reduced equipment fouling, a wider range ofupgradable feedstocks, and more efficient use of supported catalyst ifused in combination with the colloidal or molecular catalyst, ascompared to conventional hydrocracking and coking processes.

These and other advantages and features of the present invention willbecome more fully apparent from the following description and appendedclaims, or may be learned by the practice of the invention as set forthhereinafter.

BRIEF DESCRIPTION OF THE DRAWINGS

To further clarify the above and other advantages and features of thepresent invention, a more particular description of the invention willbe rendered by reference to specific embodiments thereof which areillustrated in the appended drawings. It is appreciated that thesedrawings depict only typical embodiments of the invention and aretherefore not to be considered limiting of its scope. The invention willbe described and explained with additional specificity and detailthrough the use of the accompanying drawings, in which:

FIG. 1 is a block diagram that schematically illustrates ahydroprocessing system according to the invention for upgrading a heavyoil feedstock;

FIG. 2 schematically illustrates an exemplary hydroprocessing systemthat includes a delayed coker according to the invention;

FIG. 3 is a flow diagram that schematically illustrates an exemplaryprocess for preparing a heavy oil feedstock to include a colloidal ormolecular catalyst dispersed therein;

FIG. 4 depicts a hypothetical chemical structure for an asphaltenemolecule;

FIG. 5 schematically illustrates catalyst molecules or colloidal-sizedcatalyst particles associated with asphaltene molecules;

FIGS. 6A and 6B schematically depict top and side views of a molybdenumdisulfide crystal approximately 1 nm in size;

FIG. 7 is a flow diagram that schematically illustrates an exemplaryhydroprocessing method according to the invention for upgrading a heavyoil feedstock;

FIGS. 8A and 8B are block diagrams that schematically illustratealternative exemplary embodiments of hydroprocessing systems accordingto the invention for upgrading a heavy oil feedstock;

FIG. 9 is a schematic diagram of a two-phase pre-coking hydrocrackingreactor employing a colloidal or molecular catalyst and a hot separatorupstream of a coker according to the invention;

FIG. 10 is a schematic diagram of a two-phase pre-coking hydrocrackingreactor, a three-phase ebullated bed pre-coking hydrocracking reactor,and a hot separator upstream of a coker according to the invention; and

FIG. 11 is a schematic diagram of a two-phase pre-coking hydrocrackingreactor, a hot separator, and a fixed bed pre-coking hydrocrackingreactor upstream of a coker according to the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

I. Introduction and Definitions

The present invention relates to methods and systems for upgrading heavyoil feedstock by using a hydrocracking reactor that employs a colloidalor molecular catalyst and a coking reactor downstream of thehydrocracking reactor, increasing yield and conversion rate, improvingquality of upgraded liquid hydrocarbons, and reducing formation of cokecompared to conventional hydrocracking or coking methods and systems.According to one embodiment, the methods and systems employ at least onepre-coking hydrocracking reactor, at least one separator, and at leastone coking reactor. The inventive hydroproces sing methods and systemsmay utilize the colloidal or molecular catalyst alone or in combinationwith a solid supported catalyst. The coking process downstream of thehydrocracking reactor(s) may be any coking process, such as residual oilcoking, including delayed coking, fluid coking, Flexicoking®, andcombinations or variations thereof.

The term “coking” refers to a thermal cracking process that convertspetroleum residuum, e.g., bottoms from atmospheric and vacuumdistillation of crude oil, into upgraded liquid and gas products,leaving behind a solid concentrated carbon material, or “petroleumcoke”.

The terms “coker” and “coking system” are used interchangeably and referto an apparatus system used in carrying out the coking process. The term“coking reactor” refers to an apparatus unit of a coking system in whicha substantial level of thermal cracking occurs. For example, in adelayed coking process, a coke drum is an example of a coking reactor.

The terms “coking fractionators”, “coker fractionators”, “combinationdistillation tower”, and “main fractionators” refer to the apparatusunit in a coking system that receives heavy oil feed (usually vacuumreduced crude) and coker overheads, and fractionates its contents intogases, gasoline, diesel, heavy coker gas oil, and a coke stream that isreturned to the coking reactor.

The terms “colloidal catalyst” and “colloidally-dispersed catalyst”shall refer to catalyst particles having a particle size that iscolloidal in size, e.g., less than about 500 nm in diameter, preferablyless than about 100 nm in diameter, more preferably less than about 10nm in diameter, even more preferably less than about 5 nm in diameter,and most preferably less than about 1 nm in diameter. The term“colloidal catalyst” includes, but is not limited to, molecular ormolecularly-dispersed catalyst compounds.

The terms “molecular catalyst” and “molecularly-dispersed catalyst”shall refer to catalyst compounds that are essentially “dissolved” orcompletely dissociated from other catalyst compounds or molecules in aheavy oil hydrocarbon feedstock, non-volatile liquid hydrocarbonfraction, bottoms fraction, resid, or other feedstock or product inwhich the catalyst may be found. It shall also refer to very smallcatalyst particles that only contain a few catalyst molecules joinedtogether (e.g., 15 molecules or less).

The terms “residual catalyst”, “residual molecular catalyst” and“residual colloidal catalyst” shall refer to catalyst molecules orcolloidal particles that remain with an upgraded feedstock or materialwhen transferred from one vessel to another (e.g., from a hydrocrackingreactor to a hot separator, another hydroprocessing reactor, ordistillation tower).

The term “conditioned feedstock” shall refer to a heavy oil feedstockinto which an oil soluble catalyst precursor composition has beencombined and mixed sufficiently so that, upon decomposition of thecatalyst precursor and formation of the colloidal or molecular catalystin situ, the catalyst will comprise a colloidal or molecular catalystdispersed within the feedstock.

The term “hydrocracking” shall refer to a process whose primary purposeis to reduce the boiling range of a heavy oil feedstock and in which asubstantial portion of the feedstock is converted into products withboiling ranges lower than that of the original feedstock. Hydrocrackinggenerally involves fragmentation of larger hydrocarbon molecules intosmaller molecular fragments having a fewer number of carbon atoms and ahigher hydrogen-to-carbon ratio. The mechanism by which hydrocrackingoccurs typically involves the formation of hydrocarbon free radicalsduring fragmentation followed by capping of the free radical ends ormoieties with hydrogen. The hydrogen atoms or radicals that react withhydrocarbon free radicals during hydrocracking are generated at or byactive catalyst sites.

The term “hydrotreating” shall refer to a more mild operation whoseprimary purpose is to remove impurities such as sulfur, nitrogen,oxygen, halides, and trace metals from the feedstock and saturateolefins and/or stabilize hydrocarbon free radicals by reacting them withhydrogen rather than allowing them to react with themselves. The primarypurpose is not to change the boiling range of the feedstock.Hydrotreating is most often carried out using a fixed bed reactor,although other hydroprocessing reactors can also be used forhydrotreating, an example of which is an ebullated bed hydrotreater.

Of course, “hydrocracking” may also involve the removal of sulfur andnitrogen from a feedstock as well as olefin saturation and otherreactions typically associated with “hydrotreating”. The terms“hydroprocessing” and “hydroconversion” shall broadly refer to both“hydrocracking” and “hydrotreating” processes, which define oppositeends of a spectrum, and everything in between along the spectrum.

The terms “solid supported catalyst”, “porous supported catalyst” and“supported catalyst” shall refer to catalysts that are typically used inconventional ebullated bed and fixed bed hydroprocessing systems,including catalysts designed primarily for hydrocracking orhydrodemetallization and catalysts designed primarily for hydrotreating.Such catalysts typically comprise (i) a catalyst support having a largesurface area and numerous interconnected channels or pores of unevendiameter and (ii) fine particles of an active catalyst such as sulfidesof cobalt, nickel, tungsten, and molybdenum dispersed within the pores.For example a heavy oil hydrocracking catalyst manufactured by CriterionCatalyst, Criterion 317 trilobe catalyst, has a bi-modal pore sizedistribution, with 80% of the pores ranging between 30 to 300 Angstromswith a peak at 100 Angstroms and 20% of the pores ranging between 1000to 7000 Angstroms with a peak at 4000 Angstroms. The pores for the solidcatalyst support are of limited size due to the need for the supportedcatalyst to maintain mechanical integrity to prevent excessive breakdownand formation of excessive fines in the reactor and also maintainsufficiently high surface area. Supported catalysts are commonlyproduced as cylindrical pellets, spherical solids, or extrudates.

The term “heavy oil feedstock” shall refer to heavy crude, oils sandsbitumen, bottom of the barrel and resid left over from refineryprocesses (e.g., visbreaker bottoms), and any other lower qualitymaterial that contains a substantial quantity of high boilinghydrocarbon fractions (e.g., that boil at or above 343° C. (650° F.),more particularly at or above about 524° C. (975° F.)), and/or thatinclude a significant quantity of asphaltenes that can deactivate asolid supported catalyst and/or cause or result in the formation of cokeprecursors and sediment. Examples of heavy oil feedstocks include, butare not limited to, Lloydminster heavy oil, Cold Lake bitumen, Athabascabitumen, Maya, Isthmus, Ku-Maloob-Zaap (“Ku”), Boscan, Ural, Siberian,atmospheric tower bottoms, vacuum tower bottoms, residuum (or “resid”),resid pitch, vacuum residue, and nonvolatile liquid hydrocarbonfractions that remain after subjecting crude oil, bitumen from tarsands, liquefied coal, oil shale, or coal tar feedstocks todistillation, hot separation, and the like and that contain higherboiling fractions and/or asphaltenes.

The term “hydrocracking reactor” shall refer to any vessel in whichhydrocracking (i.e., reducing the boiling range) of a feedstock in thepresence of hydrogen and a hydrocracking catalyst is the primarypurpose. Hydrocracking reactors are characterized as having an inputport into which a heavy oil feedstock and hydrogen can be introduced, anoutput port from which an upgraded feedstock or material can bewithdrawn, and sufficient thermal energy so as to form hydrocarbon freeradicals in order to cause fragmentation of larger hydrocarbon moleculesinto smaller molecules. Examples of hydrocracking reactors include, butare not limited to, two-phase hydrocracking reactors (i.e., a two phase,gas-liquid system), ebullated bed reactors (i.e., a three phase,gas-liquid-solid system), fixed bed reactors (i.e., a three-phase systemthat includes a liquid feed trickling downward over a fixed bed of solidsupported catalyst with hydrogen typically flowing cocurrently, butpossibly countercurrently in some cases).

The term “hydrocracking temperature” shall refer to a minimumtemperature required to effect significant hydrocracking of a heavy oilfeedstock. In general, hydrocracking temperatures will preferably fallwithin a range of about 395° C. (743° F.) to about 460° C. (860° F.),more preferably in a range of about 410° C. (770° F.) to about 450° C.(842° F.), and most preferably in a range of about 420° C. (788° F.) toabout 445° C. (833° F.). It will be appreciated that the temperaturerequired to effect hydrocracking may vary depending on the propertiesand chemical make-up of the heavy oil feedstock. Severity ofhydrocracking may also be imparted by varying the space velocity of thefeedstock, i.e., the residence time of feedstock in the reactor, whilemaintaining the reactor at a fixed temperature. Milder reactortemperature and longer feedstock space velocity are typically requiredfor heavy oil feedstock with high reactivity and/or high concentrationof asphaltenes.

The term “gas-liquid slurry phase hydrocracking reactor” shall refer toa hydroprocessing reactor that includes a continuous liquid phase and agaseous disperse phase which forms a “slurry” of gaseous bubbles withinthe liquid phase. The liquid phase typically comprises a hydrocarbonfeedstock that may contain a low concentration of a colloidal catalystor molecular-sized catalyst, and the gaseous phase typically compriseshydrogen gas, hydrogen sulfide, and vaporized low boiling pointhydrocarbon products. A “gas-liquid slurry phase hydrocracking reactor”should not be confused with a conventional slurry phase reactor, whichincludes three phases: a solid particulate slurry catalyst phase that istypically micron-sized or larger, a gaseous phase, and a liquid phase.

The term “gas-liquid-solid, 3-phase slurry hydrocracking reactor” isused when a solid catalyst is employed along with liquid and gas. Thegas may contain hydrogen, hydrogen sulfide and vaporized low boilinghydrocarbon products. The term “slurry phase reactor” shall broadlyrefer to both type of reactors (e.g., those with a colloidal ormolecular catalyst, those with a micron-sized or larger particulatecatalyst, and those that include both). In most cases, it shall refer toa reactor that at least includes a colloidal or molecular catalyst.

The term “asphaltene” shall refer to the fraction of a heavy oilfeedstock that is typically insoluble in paraffinic solvents such aspropane, butane, pentane, hexane, and heptane and that includes sheetsof condensed ring compounds held together by hetero atoms such assulfur, nitrogen, oxygen and metals. Asphaltenes broadly include a widerange of complex compounds having anywhere from 80 to 160,000 carbonatoms, with predominating molecular weights, as determined by solutiontechniques, in the 5000 to 10,000 range. About 80-90% of the metals inthe crude oil are contained in the asphaltene fraction which, togetherwith a higher concentration of non-metallic hetero atoms, renders theasphaltene molecules more hydrophilic and less hydrophobic than otherhydrocarbons in crude.

The terms “upgrade”, “upgrading” and “upgraded”, when used to describe afeedstock that is being or has been subjected to hydroprocessing, or aresulting material or product, shall refer to one or more of a reductionin the molecular weight of the feedstock, a reduction in the boilingpoint range of the feedstock, a reduction in the concentration ofasphaltenes, a reduction in the concentration of hydrocarbon freeradicals, and/or a reduction in the quantity of impurities, such assulfur, nitrogen, oxygen, halides, and metals.

II. Exemplary Hydroprocessing Methods and Systems

FIG. 1 schematically illustrates a hydroprocessing system 100 accordingto the invention comprising (a) a pre-coking hydrocracking reactor 104configured to receive a prepared heavy oil feedstock 102 having acolloidal or molecular catalyst dispersed therein and producing anupgraded material 106; (b) a separator 108 into which the upgradedmaterial 106 withdrawn from the pre-coking hydrocracking reactor 104 istransferred to separate gaseous and volatile fractions 116 from anon-volatile liquid hydrocarbon fraction 110; and (c) a coker 112 thatupgrades the non-volatile liquid hydrocarbon fraction 110 throughthermal cracking to form upgraded liquid hydrocarbon products 114 andpetroleum coke 118.

The prepared heavy oil feedstock 102 may comprise any desired fossilfuel feedstock and/or fraction thereof including, but not limited to,one or more of heavy crude, oil sands bitumen, bottom of the barrelfractions from crude oil, atmospheric tower bottoms, vacuum towerbottoms, coal tar, liquefied coal, and other resid fractions.

A common characteristic of heavy oil feedstocks 102 that mayadvantageously be upgraded using the hydroprocessing methods and systemsaccording to the invention is that they include a significant fractionof high boiling point hydrocarbons (i.e., at or above 343° C. (650° F.),more particularly at or above about 524° C. (975° F.)) and/orasphaltenes. The pre-coking hydrocracking reactor 104 may comprise anyhydrocracking reactor or system of reactors known in the art, includingbut not limited to, one or more gas-liquid slurry-phase reactors,gas-liquid-solid slurry-phase reactors, ebullated bed reactors, fixedbed reactors, or moving bed reactors. Ebullated bed, fixed bed, movingbed, and gas-liquid-solid slurry-phase reactors are three phase slurrysystems consisting of a solid catalyst, a liquid feedstock, and a gascontaining hydrogen, hydrogen sulfide, and vaporized low boilinghydrocarbons. A significant difference between the pre-cokinghydrocracking reactor 104 within hydroprocessing system 100 according tothe invention and conventional hydrocracking reactors is that the heavyoil feedstock 102 introduced into the pre-coking hydrocracking reactor104 includes a colloidal or molecular catalyst and/or a well-dispersedcatalyst precursor capable of forming the colloidal or molecularcatalyst in situ within feed heaters (not shown) and/or the pre-cokinghydrocracking reactor 104 itself. The colloidal or molecular catalyst,the formation of which is discussed in more detail below, may be usedalone or in combination with other catalysts, such as conventional solidsupported catalysts (e.g., porous catalysts with active catalytic siteslocated within the pores).

The separator 108 may comprise one or more hot separators, distillationtowers, fractionators, or other separators known in the art. When theseparator 108 is a hot separator, a difference between the hot separatorwithin hydroprocessing system 100 and hot separators used inconventional systems is that the upgraded feedstock or materialintroduced into the hot separator includes residual colloidal ormolecular catalyst dispersed therein as well as dissolved hydrogen. As aresult, any hydrocarbon free radicals, including asphaltene freeradicals, that are generated within the hot separator and/or whichpersist within the upgraded feedstock as withdrawn from the pre-cokinghydrocracking reactor 104 can be further hydroprocessed in the hotseparator 108.

More particularly, the colloidal or molecular catalyst within theupgraded material transferred from the pre-coking hydrocracking reactor104 to the hot separator 108 is able to catalyze beneficial upgrading orhydrotreating reactions between the hydrocarbon free radicals andhydrogen within the hot separator 108. The result is a more stableupgraded feedstock, decreased sediment and coke precursor formation, anddecreased fouling of the hot separator 108 compared to hydroprocessingsystems that do not utilize a colloidal or molecular catalyst.

The coker 112 of the hydroproces sing system 100 may be a delayed coker,fluid coker, flexicoker, or variations thereof, so long as the cokingsystem is set up to receive and process the non-volatile liquid fraction110, such as vacuum reduced crude (VRC), by thermal cracking to formupgraded liquid hydrocarbon products as the main product and petroleumcoke as byproduct. Metals contained with the liquid fraction 110,including the colloidal or molecular catalyst metal, may precipitate outwith the petroleum coke product.

The molecular or colloidal catalysts used in the system may optionallybe removed from the liquid fraction 110 before entering the one or morecoking reactors using a guard bed containing a solid supported catalyst.Thus, the hydroprocessing system in the scheme of FIG. 1 may beimplemented as comprising: (1) a heavy oil feedstock comprised of asubstantial quantity of hydrocarbons having a boiling point greater thanabout 650° F. and a colloidal or molecular catalyst dispersed throughoutthe feedstock; (2) a pre-coking hydrocracking reactor that heats ormaintains the heavy oil feedstock at a hydrocracking temperaturetogether with hydrogen during use in order to convert at least a portionof higher boiling hydrocarbons in the heavy oil feedstock to lowerboiling hydrocarbons and thereby form an upgraded material, thepre-coking hydrocracking reactor comprised of (i) an inlet port at abottom of the reactor into which the heavy oil feedstock and hydrogenare introduced and (ii) an outlet port at a top of the reactor fromwhich the upgraded material, colloidal or molecular catalyst, andhydrogen are withdrawn; (3) a separator that separates gaseous andvolatile fractions from a liquid hydrocarbon fraction in the upgradedmaterial, the separator comprised of (i) an inlet through which theupgraded material is introduced into the separator, (ii) a first outletthrough which the gaseous and volatile fractions are withdrawn, and(iii) a second outlet through which the liquid hydrocarbon fraction iswithdrawn; and (4) one or more coking reactors configured to receive andprocess the cleaned liquid hydrocarbon fraction in order to yieldupgraded hydrocarbon products and coke. An optional guard bed containinga solid supported catalyst for hydrotreating the liquid hydrocarbonfraction can be included in order for the solid supported catalyst toremove at least a portion of the residual colloidal or molecularcatalyst, metal contaminants, sulfur and other impurities from theliquid hydrocarbon fraction at some point in the process.

FIG. 2 depicts an exemplary refining system 200 that integrateshydrocracking and delayed coking according to the invention. Therefining system 200 may itself comprise a module within an even moredetailed and complex oil refinery system, including a module that isadded to a pre-existing refinery system as part of an upgrade. Therefining system 200 more particularly includes a distillation tower 202into which an initial feed 204 comprising a significant fraction ofhigher boiling hydrocarbons is introduced. By way of example and notlimitation, gases and/or lower boiling hydrocarbons 206 having a boilingpoint less than 370° C. (698° F.) are separated from a higher boilingliquid hydrocarbon fraction 208 comprising materials having a boilingpoint greater than 370° C. (698° F.). The lower boiling hydrocarbons 206can be further processed downstream either as a feed 251 to a separator232 or a stream of straight-run upgraded products 252 includinggasoline, diesel, jet fuel, kerosene, etc. In this embodiment, thehigher boiling liquid hydrocarbon fraction 208 advantageously comprisesa “heavy oil feedstock” within the meaning of this term.

An oil soluble catalyst precursor 210 is pre-blended with a hydrocarbonoil fraction or diluent 211 having a boiling range between 250-524° C.(482-975° F.), preferably having boiling point between 360-500° C.(680-932° F.), and mixed for a period of time in a pre-mixer 212 to forma diluted precursor mixture 213 in which the catalyst precursor 210 iswell-mixed with the diluent 211. By way of example, the pre-mixer 212may be an in-line static mixer. The diluted precursor mixture 213 andheavy oil feedstock 208 are combined within a conditioning chamber 214in order to thoroughly disperse the catalyst precursor 210 within theheavy oil feedstock 208 and form a conditioned feedstock 215. Theconditioning chamber 214 can be a high shear mixing apparatus and/or mayinclude a surge vessel for the feedstock 208. A surge vessel is commonlyused to dampen flow fluctuation ahead of downstream processing units andcan be used for mixing. By way of example, the mixing in chamber 214 mayconsist of a pump-around loop system.

The conditioned feedstock 215 is pressurized using a multi-stage pump218. The pump 218 may consist of many compression stages, with eachstage providing additional mixing and blending of the catalyst precursor210 within the conditioned feedstock 215 to form a finally conditionedfeedstock 216 in which the catalyst precursor 210 is more thoroughlymixed throughout the feedstock 208.

The finally conditioned feedstock 216 is introduced into a pre-heater orfurnace 220 so as to heat the finally conditioned feedstock 216 to atemperature that is about 150° C. (270° F.), preferably about 100° C.(180° F.) below the temperature in the slurry phase reactor 222. The oilsoluble catalyst precursor 210 dispersed throughout the feedstock 208decomposes and combines with sulfur released from the heavy oilfeedstock 208 to yield a colloidal or molecular catalyst as the finallyconditioned feedstock 216 travels through the pre-heater or furnace 220and is heated to a temperature higher than the decomposition temperatureof the catalyst precursor. This yields a prepared feedstock 221, whichis introduced under pressure into a slurry phase reactor 222. Hydrogengas 224, preferably preheated, is also introduced into the reactor 222under pressure to effect hydrocracking of the prepared feedstock 221within the reactor 222. Heavy oil resid bottoms 226 produced downstreamfrom the slurry phase reactor 222 may optionally be recycled back intothe reactor 222. The resid bottoms 226 may advantageously includeresidual colloidal or molecular catalyst dispersed therein. The recyclegas 228 advantageously includes hydrogen. The slurry phase reactor 222may contain a recycle channel, recycling pump, and distributor gridplate as in a conventional ebullated bed reactor to promote more evendispersion of reactants, catalyst, and heat (e.g., in a manner similarto conventional ebullated bed reactors).

The prepared feedstock 221 within the slurry phase reactor 222 is heatedor maintained at a hydrocracking temperature, which causes or allows theprepared feedstock 221, in combination with catalyst and hydrogen in thereactor 222, to be upgraded so as to form an upgraded material 230 thatis withdrawn at the top of the reactor 222. According to one embodiment,the upgraded material 230 is transferred directly to a separator 232(e.g., hot separator or distillation tower), optionally together with atleast a portion 251 of the lower boiling point fraction 206 from firstdistillation tower 202 and/or recycle gas 228 produced downstream.Alternatively, the upgraded material 230 may be introduced into one ormore hydroprocessing reactors (not shown) upstream or downstream fromthe separator 232.

Gases and volatile liquids 254 are removed from the top of 232 and sentdownstream for further processing. A higher boiling liquid hydrocarbonfraction 236 is withdrawn from the bottom of the separator 232 andfurther processed. According to one embodiment, liquid hydrocarbonfraction 236 is introduced into a vacuum distillation tower 238 in orderto separate lower boiling fractions 256 that can be volatilized throughvacuum distillation and a high boiling resid fraction 240 of lowerquality that is withdrawn from the bottom of vacuum tower 238 and eitherused as a vacuum reduced crude (VRC) 242 as a coker feed to a cokingfractionator 260 of a delayed coking system and/or as a residue 226 thatis recycled back into the slurry phase reactor 222, as discussed above.

The lower boiling fractions 256 typically include light vacuum gas oiland heavy vacuum gas oil. The lower boiling fraction 256 may be combinedwith at least a portion of the lower boiling point fraction 252withdrawn from the first distillation tower 202 and the gases andvolatile liquids 254 removed from the separator 232 and introduced intoa mixed feed hydrotreater or hydrocracking reactor (not shown) fordownstream processing to yield gasoline and other petroleum products.

The vacuum reduced crude (VRC) 242 is introduced into a cokingfractionator 260, which combines the VRC feed 242 and coker overheadfrom overhead return line 259 (hence a coking fractionator is also knownas a combination distillation tower). The coking fractionatorfractionates its contents into a lighter upgraded product stream 258 anda coking stream 261. The lighter upgraded product stream 258 comprisesgases, coker naphtha and coker gas oil.

The coking stream 261 is introduced through a coker charge pump 262 to afurnace 264. The coker charge pump 262 is normally driven by an electricmotor with a steam-driven turbine pump as a backup. The pressure may bein excess of 35 bars (500 psig) with a mechanical seal operating up to382° C. (720° F.). The coker furnace 264 heats the coking stream 261 toa coking temperature, about 500° C. (930° F.), with a pressure of about4 bars (60 psig). In an alternative embodiment, that other streams canbe added to the coker feed stream 261, such as conventional straight-runvacuum resid. For example, in some embodiments less than 100% of thecoker feed is initially processed through the slurry hydrocrackingprocess, with that balance being fed into the coker without first beinghydroprocessed by the slurry phase reactor.

Then the coking stream is introduced through a transfer line 265 and anopen feed valve 268 a to an online coke drum 270 a, wherein thermalcracking reactions occur, yielding coke and coker vapor overheads. Thetransfer line 265 can be well insulated to prevent coking and plugging.The shorter the transfer line is, the shorter the delay is betweenheating and thermal cracking reactions. This delay gives the process itsname of “delayed coking”.

Thermal cracking reactions in the coke drum 270 a cause the formation ofsolid coke and coker vapor overheads. The coker vapor overheads exitthrough coker overhead exit line 271 a and an opened coker overhead exitvalve 272 a, returning to the coking fractionator 260 through the cokeroverhead return line 259. The temperature in the exit line 271 a andreturn line 259 can be around 443° C. (830° F.). The temperature can bedecreased by about 28° C. (50° F.) by injecting hot heavy coker gas oil(not shown) into the line 271 a as quench oil to prevent coking in theline.

The delayed coking system implemented here has two alternating cokedrums to allow continuous operation, wherein the online coke drum 270 aundergoes coking reactions and the offline coke drum 270 b is undergoingdecoking. The two drums alternate between coking and decoking, allowingcontinuous operation. As illustrated here, the coke formed in theoffline coke drum 270 b is being removed by hydraulic coke cutting,wherein high pressure water is used to cut the coke out of the offlinecoke drum 270 b. Water pressures can range from 86 bars (1250 psig) to275 bars (4000 psig) and flow rates range from 2.8 cubic meters perminute (750 GPM) to 4.7 cubic meters per minute (1250 GPM). A cut waterpump 284 is a multistage barrel type or a split case multistage pump.The pump 284 can be powered with an electric motor or steam-driventurbines.

Wet coke cut off from the coke drum 270 b exits through a coke exit line273 b and an opened coke exit valve 278 b, passing through coke transferline 279 b to a wet coke receptor 280, where water is separated fromcoke and recycled through a cut water pump 284 and a water recycle line285, an open water inlet valve 276 b and water inlet line 277 b, andused for decoking again.

Decoking also involves steaming out and quenching with water. When acoke drum is undergoing decoking, the steam and hydrocarbon vapor aredirected through an opened blow-down valve 274 b and a blow-down exitline 275 to a blow-down system comprising a quench tower 290, ablow-down condenser 294 and an settling drum 296. This blow-down systemis utilized for both pollution control and for increased recovery ofhydrocarbons. During the time that a drum, such as drum 270 b is steamedout and cooled by water injection, wherein steam and hydrocarbonsstripped from the coke are directed to the quench tower 290. In quenchtower 290, hydrocarbons are condensed and returned as gas oil in a gasoil return line 291 to the coking fractionator 260. Steam mixed withhydrocarbons removed from the quench tower 290 is condensed in ablow-down condenser 294 along with an amount of oil. The oil andcondensed water are separated in the settling drum 296 and exit as gas297 a, light coker gas oil 297 b, and sour water 297 c.

The on-line off-line alteration between the two coke drums arecontrolled by a series of process stream and water valves ofcorresponding pairs, e.g., the feed valves 268 a-b, and the water inletvalves 276 a-b.

The upgraded liquid hydrocarbon products from coking 259 can beseparately or in combination with other upgraded streams 252, 254, 256further processed to produce gasoline and other upgraded hydrocarbonproducts.

A. Preparation and Characteristics of Colloidal or Molecular Catalyst

The methods according to the invention include the preliminary step ofpreparing a heavy oil feedstock so as to have a colloidal or molecularcatalyst dispersed therein, an example of which is schematicallyillustrated in the flow diagram depicted in FIG. 3 as method 300.According to one embodiment, an oil soluble catalyst precursorcomposition is pre-mixed with a diluent hydrocarbon stream to form adiluted precursor mixture, as in step 302.

The oil soluble catalyst precursor preferably has a decompositiontemperature in a range from about 100° C. (212° F.) to about 350° C.(662° F.), more preferably in a range of about 150° C. (302° F.) toabout 300° C. (572° F.), and most preferably in a range of about 175° C.(347° F.) to about 250° C. (482° F.). Examples of catalyst precursorsinclude organometallic complexes or compounds, more specifically,oil-soluble compounds or complexes of transition metals and organicacids. The catalyst precursor composition comprises at least onetransition metal and at least one organic moiety. Examples of usefultransition catalyst metals include Mo, Ni, Co, W, Fe, V and combinationsthereof. Examples of organic moieties include, but are not limited to,ligands comprising or derived from 3-cyclopentylpropionic acid,cyclohexanebutyric acid, biphenyl-2-carboxylic acid, 4-heptylbenzoicacid, 5-phenylvaleric acid, geranic acid, 10-undecenoic acid, dodecanoicacid, octanoic acid, 2-ethylhexanoic acid, naphthanic acid,pentacarbonyl, or hexacarbonyl. Exemplary precursor compositionsinclude, but are not limited to, molybdenum 2-ethylhexanoate, molybdenumnaphthanate, vanadium naphthanate, vanadium octoate, molybdenumhexacarbonyl, vanadium hexacarbonyl, iron pentacarbonyl, molybdenum3-cyclopentylpropionate, molybdenum cyclohexanebutanoate, molybdenumbiphenyl-2-carboxylate, molybdenum 4-heptylbenzoate, molybdenum5-phenylpentanoate, molybdenum geranate, molybdenum 10-undecenoate,molybdenum dodecanoate.

One of skill in the art can, following the present disclosure, select amixing temperature profile that results in intimate mixing of a selectedprecursor composition without substantial decomposition prior toformation of the colloidal or molecular catalyst.

Examples of suitable hydrocarbon diluents include, but are not limitedto, vacuum gas oil (which typically has a boiling range of 360-524° C.)(680-975° F.), decant oil or cycle oil (which typically has a boilingrange of 360°-550° C.) (680-1022° F.), and light gas oil (whichtypically has a boiling range of 200°-360° C.) (392-680° F.).

The ratio of catalyst precursor to hydrocarbon oil diluent is preferablyin a range of about 1:500 to about 1:1, more preferably in a range ofabout 1:150 to about 1:2, and most preferably in a range of about 1:100to about 1:5 (e.g., 1:100, 1:50, 1:30, or 1:10).

The catalyst precursor is advantageously mixed with the hydrocarbondiluent at a temperature below which a significant portion of thecatalyst precursor starts to decompose, preferably, at temperature in arange of about 25° C. (77° F.) to about 250° C. (482° F.), morepreferably in a range of about 50° C. (122° F.) to about 200° C. (392°F.), and most preferably in a range of about 75° C. (167° F.) to about150° C. (302° F.), to form the diluted precursor mixture. It will beappreciated that the actual temperature at which the diluted precursormixture is formed typically depends largely on the decompositiontemperature of the particular precursor that is utilized. The catalystprecursor is preferably mixed with the hydrocarbon oil diluent for atime period in a range of about 1 second to about 20 minutes, morepreferably in a range of about 5 seconds to about 10 minutes, and mostpreferably in a range of about 20 seconds to about 5 minutes. The actualmixing time is dependent, at least in part, on the temperature (i.e.,which affects the viscosity of the fluids) and mixing intensity. Mixingintensity is dependent, at least in part, on the number of stages e.g.,for in-line static mixers.

Whereas it is within the scope of the invention to directly blend thecatalyst precursor with heavy oil feedstocks, care must be taken in suchcases to mix the components for a time sufficient to thoroughly blendthe precursor within the feedstock before substantial decomposition ofthe precursor has occurred. For example, U.S. Pat. No. 5,578,197 to Cyret al., the disclosure of which is incorporated by reference, describesa method whereby molybdenum 2-ethylhexanoate is mixed with bitumenvacuum tower residuum for 24 hours before the resulting mixture washeated in a reaction vessel to form the catalyst compound and to effecthydrocracking (see col. 10, lines 4-43). Whereas 24-hour mixing in atesting environment may be acceptable, such long mixing times may makecertain industrial operations prohibitively expensive.

Pre-blending the catalyst precursor with a hydrocarbon diluent prior toblending the diluted precursor mixture with the heavy oil feedstockgreatly aids in thoroughly and intimately blending the precursor withinthe feedstock, particularly in the relatively short period of timerequired for large-scale industrial operations to be economicallyviable. Forming a diluted precursor mixture shortens the overall mixingtime by (1) reducing or eliminating differences in solubility betweenthe more polar catalyst precursor and the heavy oil feedstock, (2)reducing or eliminating differences in rheology between the catalystprecursor and the heavy oil feedstock, and/or (3) breaking up thecatalyst precursor molecules to form a solute within a hydrocarbon oildiluent that is much more easily dispersed within the heavy oilfeedstock. It is particularly advantageous to first form a dilutedprecursor mixture in the case where the heavy oil feedstock containswater (e.g., condensed water). Otherwise, the greater affinity of thewater for the polar catalyst precursor can cause localized agglomerationof the precursor, resulting in poor dispersion and formation ofmicron-sized or larger catalyst particles. The hydrocarbon oil diluentis preferably substantially water free (i.e., contains less than about0.5% water) to prevent the formation of substantial quantities ofmicron-sized or larger catalyst particles.

The diluted precursor mixture is combined with a heavy oil feedstock, asin step 304 of method 300 illustrated in FIG. 3, and mixed for a timesufficient and in a manner so as to disperse the catalyst precursorthroughout the feedstock and yield a conditioned feedstock in which thecatalyst precursor is thoroughly mixed within the heavy oil feedstockprior to precursor decomposition. In order to obtain sufficient mixingof the catalyst precursor within the heavy oil feedstock to thereafteryield a colloidal or molecular catalyst upon decomposition of theprecursor, the diluted precursor mixture and heavy oil feedstock arepreferably mixed for a time period in a range of about 1 second to about20 minutes, more preferably in a range from about 5 second to about 10minutes, and most preferably in a range of about 20 seconds to about 3minutes. Increasing the vigorousness and/or shearing energy of themixing process generally reduce the time required to effect thoroughmixing.

Examples of mixing apparatus that can be used to effect thorough mixingof the catalyst precursor and heavy oil feedstock include, but are notlimited to, high shear mixing such as mixing created in a vessel with apropeller or turbine impeller; multiple static in-line mixers; multiplestatic in-line mixers in combination with in-line high shear mixers;multiple static in-line mixers in combination with in-line high shearmixers; multiple static in-line mixers in combination with in-line highshear mixers followed by a pump around in the surge vessel; combinationsof the above followed by one or more multi-stage centrifugal pumps; andone or more multi-stage centrifugal pumps. According to one embodiment,continuous rather than batch-wise mixing can be carried out using highenergy pumps having multiple chambers within which the catalystprecursor composition and heavy oil feedstock are churned and mixed aspart of the pumping process itself. The foregoing mixing apparatus mayalso be used for the pre-mixing process discussed above in which thecatalyst precursor is mixed with the hydrocarbon oil diluent to form thecatalyst precursor mixture.

According to one embodiment, the diluted precursor mixture can beinitially mixed with about 10-30% (e.g., 20%) of the heavy oilfeedstock, the resulting mixed heavy oil feedstock can be mixed in withanother 30-50% (e.g., 40%) of the heavy oil feedstock, and the resulting50-70 (e.g., 60%) of the mixed heavy oil feedstock can be mixed in withthe remainder 30-50% (e.g., 40%) of heavy oil in accordance with goodengineering practice of progressive dilution to thoroughly disperse thecatalyst precursor in the heavy oil feedstock. The foregoing percentagesare merely illustrative and non-limiting. Vigorous adherence to themixing time in the appropriate mixing devices or methods describedherein should still be used in the progressive dilution approach.

In the case of heavy oil feedstocks that are solid or extremely viscousat room temperature, such feedstocks may advantageously be heated inorder to soften them and create a feedstock having sufficiently lowviscosity so as to allow good mixing of the catalyst precursor into thefeedstock composition. These heavy oil feedstocks include, but are notlimited to heavy crude oil, oil sand bitumen, atmospheric tower bottoms,vacuum tower bottoms, resid, visbreaker bottoms, coal tar, heavy oilfrom oil shale, and liquefied coal. In general, decreasing the viscosityof the heavy oil feedstock will reduce the time required to effectthorough and intimate mixing of the oil-soluble precursor compositionwithin the feedstock. However, the feedstock should not be heated to atemperature above which significant decomposition of the catalystprecursor occurs until after thorough and complete mixing to form theconditioned feedstock. Premature decomposition of the catalyst precursorgenerally results in formation of micron-sized or larger catalystparticles rather than a colloidal or molecular catalyst. The heavy oilfeedstock and diluted precursor mixture are preferably mixed andconditioned at a temperature in a range of about 25° C. (77° F.) toabout 350° C. (662° F.), more preferably in a range of about 50° C.(122° F.) to about 300° C. (572° F.), and most preferably in a range ofabout 75° C. (167° F.) to about 250° C. (482° F.) to yield theconditioned feedstock.

After the catalyst precursor has been well-mixed throughout the heavyoil feedstock to yield a conditioned feedstock, this composition isheated to a temperature where significant decomposition of the catalystprecursor occurs in order to liberate the catalyst metal therefrom toform the final active catalyst. According to one embodiment, metal fromthe precursor reacts with sulfur liberated from the heavy oil feedstockto yield a metal sulfide compound that is the final active catalyst.Where the heavy oil feedstock includes sufficient or excess sulfur, thefinal activated catalyst may be formed in situ by heating to atemperature sufficient to liberate the sulfur therefrom. In some cases,sulfur may be liberated at the same temperature that the precursordecomposes. In other cases, further heating to a higher temperature maybe required.

If the catalyst precursor is thoroughly mixed throughout the heavy oilfeedstock, at least a substantial portion of the liberated metal ionswill be sufficiently sheltered or shielded from other metal ions so thatthey can form a molecularly-dispersed catalyst upon reacting with sulfurto form the metal sulfide compound. Under some circumstances, minoragglomeration may occur, yielding colloidal-sized catalyst particles.However, it is believed that taking care to thoroughly mix the precursorthroughout the feedstock will yield individual catalyst molecules ratherthan particles. Simply blending, while failing to sufficiently mix, thecatalyst precursor with the feedstock typically causes formation oflarge agglomerated metal sulfide compounds that are micron-sized orlarger.

In order to form the metal sulfide catalyst, the blended feedstockcomposition is preferably heated to a temperature in a range of about225° C. (437° F.) to about 450° C. (842° F.), more preferably in a rangeof about 275° C. (527° F.) to about 440° C. (824° F.), and mostpreferably in a range of about 310° C. (590° F.) to about 420° C. (788°F.). According to one embodiment, the conditioned feedstock is heated toa temperature that is about 150° C. (270° F.) less than thehydrocracking temperature within the pre-coking hydrocracking reactor,preferably about 100° C. (180° F.) less than the hydrocrackingtemperature. According to one embodiment, the colloidal or molecularcatalyst is formed during preheating before the heavy oil feedstock isintroduced into the pre-coking hydrocracking reactor. According toanother embodiment, at least a portion of the colloidal or molecularcatalyst is formed in situ within the pre-coking hydrocracking reactoritself. In some cases, the colloidal or molecular catalyst can be formedas the heavy oil feedstock is heated to a hydrocracking temperatureprior to or after being introduced into a hydrocracking reactor. Theinitial concentration of catalyst metal in the colloidal or molecularcatalyst is preferably in a range of about 5 parts per million (ppm) toabout 500 ppm by weight of the heavy oil feedstock, more preferably in arange of about 15 ppm to about 300 ppm, and most preferably in a rangeof about 25 ppm to about 175 ppm. The catalyst may become moreconcentrated (e.g., up to 5000 ppm) as volatile fractions are removedfrom a non-volatile resid fraction.

In the case where the heavy oil feedstock includes a significantquantity of asphaltene molecules, the catalyst molecules or colloidalparticles will preferentially associate with, or remain in closeproximity to, the asphaltene molecules. For one embodiment of theinvention, the heavy oil feedstock comprises at least about 10 wt %asphaltenes or other coke forming precursors. The upgrading reactions inthe process according to the invention reduce the quantity ofasphaltenes or other coke forming precursors by at least 20 wt %, morepreferably by at least 40 wt %, and even more preferably by at least 60wt %.

Asphaltenes are complex hydrocarbon molecules that include a relativelylow ratio of hydrogen to carbon that is the result of a substantialnumber of condensed aromatic and naphthenic rings with paraffinic sidechains. A hypothetical asphaltene molecule structure developed by A.G.Bridge and co-workers at Chevron is depicted in FIG. 4. Sheetsconsisting of the condensed aromatic and naphthenic rings are heldtogether by heteroatoms such as sulfur or nitrogen and/or polymethylenebridges, thio-ether bonds, and vanadium and nickel complexes. Theasphaltene fraction also contains a higher content of sulfur andnitrogen than does crude oil or the rest of the vacuum resid, and italso contains higher concentrations of carbon-forming compounds (i.e.,that form coke precursors and sediment).

Asphaltene has a greater affinity for the colloidal or molecularcatalyst since asphaltene molecules are generally more hydrophilic andless hydrophobic than other hydrocarbons contained within the heavy oilfeedstock. Because the colloidal or molecular catalyst tends to be veryhydrophilic, the individual particles or molecules will tend to migratetoward the more hydrophilic moieties or molecules within the heavy oilfeedstock. FIG. 5 schematically depicts catalyst molecules, or colloidalparticles “X” associated with, or in close proximity to, the asphaltenemolecules.

While the highly polar nature of the catalyst compound causes or allowsthe colloidal or the molecular catalyst to associate with asphaltenemolecules, it is the general incompatibility between the highly polarcatalyst compound and the hydrophobic heavy oil feedstock thatnecessitates the aforementioned intimate or thorough mixing of the oilsoluble catalyst precursor composition within the heavy oil feedstockprior to decomposition of the precursor and formation of the colloidalor molecular catalyst. Because metal catalyst compounds are highlypolar, they cannot be effectively dispersed within a heavy oil feedstockin colloidal or molecular form if added directly thereto or as part ofan aqueous solution or an oil and water emulsion. Such methodsinevitably yield micron-sized or larger catalyst particles.

Reference is now made to FIGS. 6A and 6B, which schematically depict ananometer-sized molybdenum disulfide crystal. FIG. 6A is a top view, andFIG. 6B is a side view of a molybdenum disulfide crystal. Molecules ofmolybdenum disulfide typically form flat, hexagonal crystals in whichsingle layers of molybdenum (Mo) atoms are sandwiched between layers ofsulfur (S) atoms. The only active sites for catalysis are on the crystaledges where the molybdenum atoms are exposed. Smaller crystals have ahigher percentage of molybdenum atoms exposed at the edges.

The diameter of a molybdenum atom is approximately 0.3 nm, and thediameter of a sulfur atom is approximately 0.2 nm. A nanometer-sizedcrystal of molybdenum disulfide has 7 molybdenum atoms sandwiched inbetween 14 sulfur atoms. As seen in FIG. 6A, 6 out of 7 (85.7%) of thetotal molybdenum atoms will be exposed at the edge and available forcatalytic activity. In contrast, a micron-sized crystal of molybdenumdisulfide has several million atoms, with only about 0.2% of the totalmolybdenum atoms being exposed at the crystal edge and available forcatalytic activity. The remaining 99.8% of the molybdenum atoms in themicron-sized crystal are embedded within the crystal interior and aretherefore unavailable for catalysis. This means that nanometer-sizedmolybdenum disulfide particles are, at least in theory, orders ofmagnitude more efficient than micron-sized particles in providing activecatalyst sites.

In practical terms, forming smaller catalyst particles results in morecatalyst particles and more evenly distributed catalyst sites throughoutthe feedstock. Simple mathematics dictates that forming nanometer-sizedparticles instead of micron-sized particles will result in approximately100³ (or 1 million) to 1000³ (or 1 billion) times more particlesdepending on the size and shape of the catalyst crystals. That meansthere are approximately 1 million to 1 billion times more points orlocations within the feedstock where active catalyst sites reside.Moreover, nanometer-sized or smaller molybdenum disulfide particles arebelieved to become intimately associated with asphaltene molecules, asshown in FIG. 5. In contrast, micron-sized or larger catalyst particlesare believed to be far too large to become intimately associated with orwithin asphaltene molecules.

B. Hydroprocessing Methods

FIG. 7 is a flow diagram that schematically illustrates an examplemethod 700 for hydroproces sing a heavy oil feedstock according to theinvention to form increased quantity and quality of upgraded liquidhydrocarbon products and reduced quantity of coke, as compared toconventional hydrocracking or thermal cracking upgrading methods.

First, a heavy oil feedstock is introduced together with hydrogen into apre-coking hydrocracking reactor, as in step 702 of method 700. Theheavy oil feedstock initially comprising at least about 30 wt %, or morepreferably at least about 50 wt %, or even more preferably at leastabout 80 wt %, of hydrocarbons having a boiling point of at least about524° C. (975° F.). The colloidal or molecular catalyst may be formed insitu within the heavy oil feedstock prior to introducing the feedstockin the pre-coking hydrocracking reactor, or at least a portion of thecolloidal or molecular catalyst may be generated in situ within thepre-coking hydrocracking reactor itself. Examples of suitablehydrocracking reactors that may be used in this first step or subpart ofthe method include gas-liquid slurry phase reactors, gas-liquid-solidslurry phase reactors, ebullated bed reactors, fixed bed reactors andmoving bed reactors.

Then, the heavy oil feedstock is heated to or maintained at ahydrocracking temperature so as to form hydrocarbon free radicals fromhydrocarbon molecules in the heavy oil feedstock, as in step 704 ofmethod 700. The feedstock may be introduced into the pre-cokinghydrocracking reactor already heated to the hydrocracking temperature,or may be heated within the pre-coking hydrocracking reactor to thehydrocracking temperature in order to yield the hydrocarbon freeradicals.

The colloidal or molecular catalyst within the feedstock catalyzesupgrading reactions between the hydrocarbon free radicals and thehydrogen within the pre-coking hydrocracking reactor to form an upgradedfeedstock or material, as in step 706 of method 700. The upgradingreactions reduce the quantity of asphaltenes or other coke formingprecursors, increase the hydrogen to carbon ratio in the upgradedmaterial, and decrease the boiling points of hydrocarbons in theupgraded material compared to the heavy oil feedstock. According to oneembodiment, excess hydrogen is introduced into the pre-cokinghydrocracking reactor in order to ensure high conversion levels andthroughput.

The upgraded material is withdrawn from the pre-coking hydrocrackingreactor and transferred to a separator, as in step 708 of method 700. Asdiscussed in more detail below, the upgraded material may alternativelybe introduced into one or more additional pre-coking hydroprocessingreactors for further upgrading upstream or downstream from theseparator. The separator allows or causes separation of gases andvolatile liquids from a non-volatile liquid hydrocarbon fraction, as instep 710 of method 700. The gaseous and volatile liquid hydrocarbonfractions are advantageously withdrawn from the top of the separator andthe non-volatile (or less volatile) liquid hydrocarbon fraction iswithdrawn from the bottom of the separator.

In one embodiment, the separator comprises a hot separator. Because ofthe high temperature in the hot separator, residual colloidal ormolecular catalyst within the liquid hydrocarbon fraction may continueto catalyze hydrogenation reactions between residual or newly addedhydrogen and hydrocarbon free radicals that persist and/or that arenewly generated within the hot separator. Because catalytichydrogenation reactions continue, the temperature within the hotseparator can be maintained at or near the hydrocracking temperaturewithout fouling the equipment as a result of the formation of cokeprecursors and sediment, which could otherwise occur within the hotseparator if the asphaltene radicals were allowed to react with eachother and other hydrocarbon radicals instead of being capped withhydrogen.

At least a portion of the non-volatile (or less volatile) liquidhydrocarbon fraction withdrawn from the bottom of the separator istransferred to one or more coking reactors, as in step 712 of method700, causing thermal cracking of the liquid hydrocarbon fraction to formupgraded hydrocarbon products and petroleum coke. In an implementationof the invention, at least a portion of the liquid hydrocarbon fractionintroduced into the coking reactor comprises vacuum reduced crude (VRC).The coking reactors may be delayed coking reactors, fluid cokingreactors, Flexicoking® reactors, or other coking reactors.

The molecular or colloidal catalysts used in the process may berecovered from the process stream as a recycle at the bottom of aseparator or a distillation column. Optionally, the hydroprocessingmethod comprises introducing the liquid hydrocarbon fraction into aguard bed containing a solid supported catalyst before introducing itinto the one or more coking reactors, the solid supported catalystremoving at least a portion of the residual colloidal or molecularcatalyst and impurities in the liquid process stream.

Finally, the upgraded hydrocarbon products and coke are separated forfurther processing or storage, as in step 714 of method 700. Thepre-coking hydrocracking reactions facilitated by the colloidal ormolecular catalyst reduce coke-forming asphaltenes and other precursors,and increase hydrogen to carbon ratio of the processed stream. Theprocess according to the invention leads to higher conversion level andyield and improved quality of upgraded hydrocarbons, as compared toconventional upgrading processes. It also reduces coke formation andequipment fouling, even in the coker, enabling processing of a widerrange of lower quality feedstocks.

For the purpose of this invention, the term “conversion” is definedbased on the amount of coke formed in the coker. In other words, %Conversion=100×(Resid Feedstock−Coke Formed)/(Resid Feedstock). Thedisclosed methods typically convert at least about 60 wt % ofhydrocarbons from having a boiling point of at least 524° C. (975° F.)to having one below 524° C. (975° F.); preferably at least about 70 wt%, more preferably at least about 80 wt %, and most preferably at leastabout 85 wt %. An embodiment of the invention improves conversion rateby at least 4 wt %, or more preferably by at least 10 wt %, compared toan otherwise analogous coking method in the absence of pre-cokinghydrocracking reactions catalyzed by the molecular or colloidalcatalyst.

A method according to the invention yields at least 70 wt % of C₄+hydrocarbons and a boiling point of less than 524° C. (975° F.).Preferably, the method improves yield of C₄+ hydrocarbons and a boilingpoint of less than 524° C. (975° F.) by at least 10% compared to anotherwise analogous coking method in the absence of pre-cokinghydrocracking reactions catalyzed by the molecular or colloidalcatalyst. A method implemented according to the invention reduces cokeformation by at least 20 wt %, or more preferably by at least 50 wt %,compared to an otherwise analogous coking method without in the absenceof pre-coking hydrocracking reactions catalyzed by the molecular orcolloidal catalyst.

C. Hydroprocessing Systems

As mentioned above, the processing stream of hydrocarbons may undergoadditional upgrading reactions before being processed by a cokingreactor. FIGS. 8A-8B schematically illustrates exemplary systemconfigurations in accordance with the invention that are variations ofthe configuration shown in FIG. 1. The system of FIG. 1 can beimplemented to include a slurry phase pre-coking hydrocracking reactor902 and a hot separator 904, the design and operation of which are shownand described more fully below with respect to FIG. 9. In an alternativeembodiment, at least one pre-coking hydrocracking reactor may comprisean ebullated bed hydrocracking reactor 1030, the design and operation ofwhich are shown and described more fully below with respect to FIG. 10.In another embodiment, at least one pre-coking hydrocracking reactor maycomprise a fixed bed hydrocracking reactor 1160, the design andoperation of which are shown and described more fully below with respectto FIG. 11.

Variations from the system above may have configurations exemplified inFIGS. 8A-8B. A system configuration may include a second pre-cokinghydrocracking reactor and optionally a second separator. Each of the twohydrocracking reactors may comprise, but are not limited to, a slurryphase reactor, an ebullated bed reactor, or a fixed bed reactor. Thefirst of the two pre-coking hydrocracking reactors is preferably aslurry phase reactor that includes a liquid phase comprising the heavyoil feedstock and colloidal or molecular catalyst dispersed therein anda gaseous phase comprising mainly hydrogen. The separators may compriseone or more hot separators, distillation towers, fractionators, or anyother separators known in the art, wherein the first of the twoseparators is preferably a hot separator. Additional variations ofconfigurations based on the configurations of FIG. 1 can be devised byone skilled in the art according to the invention—combininghydrocracking reactions involving a molecular/colloidal catalyst withthermal cracking reactions of coking.

The embodiment shown in FIG. 8A adds a second pre-coking hydrocrackingreactor 808 in the system after the first pre-coking hydrocrackingreactor 804 and before the separator 812. The system comprises at least(1) a heavy oil feedstock 802 comprised of a substantial quantity ofhydrocarbons having a boiling point greater than about 650° F. and acolloidal or molecular catalyst dispersed throughout the feedstock; (2)a pre-coking hydrocracking reactor 804 that heats or maintains the heavyoil feedstock 802 at a hydrocracking temperature together with hydrogenin order to convert at least a portion of higher boiling hydrocarbons inthe heavy oil feedstock to lower boiling hydrocarbons and thereby forman upgraded material 806; (3) a second pre-coking hydrocracking reactor808 that further upgrades the upgraded material 806 to form a furtherupgraded material 810 to be introduced into a separator 812; (4) theseparator 812 separates gaseous and volatile fractions 816 from a liquidhydrocarbon fraction 814 in the further upgraded material 810; and (5)one or more coking systems 818 configured to receive and process theliquid hydrocarbon fraction 814 and to form and separate upgraded liquidhydrocarbon products 820 and coke 822.

The alternative embodiment shown in FIG. 8B adds a second pre-cokinghydrocracking reactor 854 and a second separator 858 in the systembetween the first separator 848 and the coker 868. The second pre-cokinghydrocracking reactor 858 upgrades the liquid hydrocarbon fraction 850from the first separator 848 to form a further upgraded material 856 tobe introduced into the second separator 858, which separates gaseous andvolatile fractions 862 from a liquid hydrocarbon fraction 860 in thefurther upgraded material 856, which is then processed by the coker 868,yielding upgraded liquid hydrocarbon products 870 and coke 872.

FIG. 9 schematically depicts a hydroprocessing system 900 that includesa two-phase slurry phase hydrocracking reactor 902 and a hot separator904. A heavy oil feedstock 906 is blended and conditioned with acatalyst precursor 908 within a mixer 910, preferably after firstpre-mixing the precursor composition 908 with a diluent 907 to formdiluted precursor mixture 909. The conditioned feedstock from the mixer910 is pressurized by a pump 912, which also serves as a multi-stagemixing device to further disperse the catalyst precursor composition asdiscussed above, passed through a pre-heater 914, and continuously fedinto the reactor 902 together with hydrogen gas 916 through an inputport 918 located at or near a bottom of the reactor 902. A stirrer 920at the bottom of the reactor 902 induces mixing within the liquid phase,thus helping to more evenly disperse the heat generated by thehydrocracking reactions. Alternatively, or in addition to the stirrer920, the reactor 902 may include a recycle channel, recycling pump, anddistributor grid plate (not shown) as in conventional ebullated bedreactors (See FIG. 10) to promote more even dispersion of reactants,catalyst, and heat. Hydrogen is schematically depicted as gas bubbles922, within the feedstock 306. The colloidal or molecular catalystwithin the feedstock 906 is schematically depicted as catalyst particles924. It will be appreciated that gas bubbles 922 and catalyst particles924 are shown oversized so that they may be seen in the drawing.

The heavy oil feedstock 906 is catalytically upgraded in the presence ofthe hydrogen and colloidal or molecular catalyst within the slurry phasereactor 902 to form an upgraded feedstock 926, which is continuouslywithdrawn from the slurry phase reactor 902 through an output port 928located at or near the top of the slurry phase reactor 902 and then fedinto the separator 904, optionally after passing through optionalhydroprocessing apparatus 930.

The upgraded feedstock 926 in hot separator 904 contains residual ormolecular catalyst, schematically depicted as catalyst particles 924′within the hot separator 904, and residual hydrogen, schematicallydepicted as bubbles 922′. The hot separator 904, which may be operatedat a temperature within about 20° F. (about 11° C.) of the hydrocrackingtemperature within the reactor 902, separates the volatile fraction 905,which is withdrawn from the top of hot separator 904, from thenon-volatile liquid fraction 907, which is withdrawn from the bottom ofhot separator 904. Beneficial upgrading reactions between hydrocarbonfree radicals that still exist and/or are formed within non-volatilefraction 907 and residual hydrogen 922′ can be catalyzed by residualcolloidal or molecular catalyst 924′ within the hot separator 904.

The liquid fraction 907 is withdrawn from the hot separator 904 andtransferred to one or more coking reactors to form upgraded liquidhydrocarbon products and coke. The colloidal or molecular catalysts inthe non-volatile fraction 907 may optionally be removed by a guard bed(not shown) containing a solid supported catalyst before introducingliquid fraction 907 into one or more coking reactors.

FIG. 10 schematically depicts a hydroprocessing system 1000 thatincludes a slurry phase hydrocracking reactor 1002, a hot separator1004, and an ebullated bed reactor 1030 disposed between the slurryphase reactor 1002 and hot separator 1004. The slurry phase reactor 1002produces an upgraded feedstock 1026 in essentially the same way as inhydroprocessing system 900, except that the upgraded feedstock 1026 isfed into the ebullated bed reactor 1030 instead of the hot separator904. The upgraded feedstock 1026 is optionally pressurized by pump 1032and introduced together with supplemental hydrogen 1034 into theebullated bed reactor 1030 through an input port 1036 located at or nearthe bottom of the ebullated bed reactor 1030. The upgraded feedstock1026 contains residual or molecular catalyst, schematically depicted ascatalyst particles 1024′ within the ebullated bed reactor 1030. Theebullated bed reactor 1030 also includes an output port 1038 at or nearthe top of the ebullated bed reactor 1030 through which a furtherhydroprocessed feedstock 1040 is withdrawn.

The ebullated bed reactor 1030 further includes an expanded catalystzone 1042 comprising a porous supported catalyst 1044 that is maintainedin an expanded or fluidized state against the force of gravity by upwardmovement of feedstock and gas through the ebullated bed reactor 1030.The lower end of the expanded catalyst zone 1042 is defined by adistributor grid plate 1046 with bubble caps, which separates theexpanded catalyst zone 1042 from a lower supported catalyst free zone1048 located between the bottom of the ebullated bed reactor 1030 andthe distributor grid plate 1046. The distributor grid plate 1046distributes the hydrogen gas and feedstock evenly across the reactor andprevents the supported catalyst 1044 from falling by the force ofgravity into the lower supported catalyst free zone 1048. The upper endof the expanded catalyst zone 1042 is the height at which the downwardforce of gravity begins to equal or exceed the uplifting force of theupwardly moving feedstock and gas through the ebullated bed reactor 1030as the supported catalyst 1044 reaches a given level of expansion orseparation. Above the expanded catalyst zone 1042 is an upper supportedcatalyst free zone 1050. Residual colloidal or molecular catalyst 1024′is dispersed throughout the feedstock within the ebullated bed reactor1030, including both the expanded catalyst zone 1022 and the supportedcatalyst free zones 1048, 1050.

Feedstock within the ebullated bed reactor 1030 continuouslyrecirculates from the upper supported catalyst free zone 1050 to thelower supported catalyst free zone 1048 of the ebullated bed reactor1030 by means of a recycling channel 1052 disposed in the center of theebullated bed reactor 1030 in communication with an ebullating pump 1054disposed at the bottom of the ebullated bed reactor 1030. At the top ofthe recycling channel 1052 is a funnel-shaped recycle cup 1056 throughwhich feedstock is drawn from the upper supported catalyst free zone1050. The feedstock drawn downward through the recycling channel 1052enters the lower catalyst free zone 1048 and then passes up through thedistributor grid plate 1046 and into the expanded catalyst zone 1042,where it is blended with fresh upgraded feedstock 1026 and supplementalhydrogen gas 1034 entering the ebullated bed reactor 1030 through theinput port 1036. Continuously circulating blended feedstock upwardthrough the ebullated bed reactor 1030 advantageously maintains thesupported catalyst 1044 in an expanded or fluidized state within theexpanded catalyst zone 1042, minimizes channeling, controls reactionrates, and keeps heat released by the exothermic hydrogenation reactionsto a safe level.

Fresh supported catalyst 1044 is introduced into the ebullated bedreactor 1030, more specifically the expanded catalyst zone 1042, througha catalyst input tube 1058 that passes through the top or bottom of theebullated bed reactor 1030 and directly into the expanded catalyst zone1042. Spent supported catalyst 1044 is withdrawn from the expandedcatalyst zone 1042 through a catalyst withdrawal tube 1060 that passesfrom a lower end of the expanded catalyst zone 1042 through both thedistributor grid plate 1046 and the bottom of the ebullated bed reactor1030. It will be appreciated that the catalyst withdrawal tube 1060 isunable to differentiate between fully spent catalyst, partially spentbut active catalyst, and fresh catalyst such that a random distributionof supported catalyst 1044 is withdrawn from the ebullated bed reactor1030 as “spent” catalyst. This has the effect of wasting a certainamount of the supported catalyst 1044. On the other hand, the existenceof residual colloidal or molecule catalyst, schematically shown ascatalyst particles 1024′, within the ebullated bed reactor 1030,provides additional catalytic hydrogenation activity, both within theexpanded catalyst zone 1042 and the lower and upper supported catalystfree zones 1048, 1050. Capping of free radicals minimizes formation ofsediment and coke precursors, which are often responsible fordeactivating the supported catalyst. This may have the effect ofreducing the amount of supported catalyst 1044 that would otherwise berequired to carry out a desired hydroprocessing reaction. It may alsoreduces the rate at which the supported catalyst 1044 must be withdrawnand replenished.

Then, the further hydroprocessed feedstock 1040 withdrawn from theebullated bed reactor 1030 is introduced into the hot separator 1004.The hot separator 1004, which is advantageously operated at atemperature within about 20° F. (about 11° C.) of the hydroprocessingtemperature within the ebullated bed reactor 1030, separates thevolatile fraction 1005, which is withdrawn from the top of hot separator1004, from the non-volatile fraction 1007, which is withdrawn from thebottom of hot separator 1004. The non-volatile fraction 1007 typicallycontains residual colloidal or molecular catalyst, schematicallydepicted as catalyst particles 1024″, and residual hydrogen gas,schematically depicted as bubbles 1022″, dispersed therein.

Finally, the non-volatile fraction 1007 can be withdrawn from the hotseparator 1004 and transferred to one or more coking reactors to formupgraded liquid hydrocarbon products and coke.

FIG. 11 schematically depicts a hydroproces sing system 1100 thatincludes a slurry phase reactor 1102, a hot separator 1104, and a fixedbed reactor 1160. The slurry phase reactor 1102 produces an upgradedfeedstock 1118 in essentially the same way as the slurry phase reactor902 in hydroprocessing system 900 (FIG. 9), and the hot separator 1104separates a volatile fraction 1105 from a non-volatile fraction 1107 inessentially the same way as the hot separator 904 in hydroprocessingsystem 900. The upgraded feedstock 1118 contains residual colloidal ormolecular catalyst, schematically depicted as catalyst particles 1124′within the hot separator 1104. The non-volatile fraction 1107 is,however, introduced into the fixed bed reactor 1160 for furtherhydroprocessing. The fixed bed reactor 1160 may be designed to performhydrocracking and/or hydrotreating reactions depending on the operatingtemperature and/or the type of solid supported catalyst that is usedwithin the fixed bed reactor 1160.

Fixed bed reactor 1160 more particularly includes an input port 1162 atthe top through which the non-volatile fraction 1107 and supplementalhydrogen gas 1164 are introduced, and an output port 1166 at the bottomthrough which a further hydroprocessed feedstock 1188 is withdrawn. Thefixed bed reactor 1160 further includes a plurality of verticallystacked and spaced apart catalyst beds 1170 comprising a packed poroussupported catalyst. Above each catalyst bed 1170 is a distributor grid1172, which helps to more evenly distribute the flow of feedstockdownward through the catalyst beds 1170. Supported catalyst free zones1174 exist above and below each catalyst bed 1170. To the extent theresidual colloidal or molecular catalyst 1124′ is not preliminarilyremoved by a guard bed, it remains dispersed throughout the feedstockwithin the fixed bed reactor 1160, in both the catalyst beds 1170 andthe supported catalyst free zones 1174. Auxiliary ports 1176 in thecenter and/or bottom of the fixed bed reactor 1160 may be providedthrough which a cooling oil and/or hydrogen quench can be introduced tocool heat generated by the hydroprocessing reactions, control thereaction rate, and thereby help prevent formation of coke precursors andsediment and/or excessive gas within the fixed bed reactor 1160.

The further hydroprocessed feedstock 1188 is then withdrawn from thefixed bed reactor 1160, separated into volatile and nonvolatilefractions, the latter of which is transferred to one or more cokingreactors that yields upgraded hydrocarbon products and coke, which areseparated by the coking system for storage and further downstreamprocessing.

III. Examples

The following examples describe test studies which demonstrate theeffects and advantages of hydrocracking petroleum atmospheric or vacuumresidues in a hydrocracking reactor using a colloidal or molecularcatalyst followed by processing of the resultant residue in a coker, ascompared to a conventional hydrocracking process or a conventionalcoking process.

The heavy oil feedstock used in the examples include: blend of 75 W %Cold Lake and 25 W % Athabasca vacuum residues, Black Rock atmosphericresidue, Black Rock vacuum residue, Murphy vacuum residue, Ku vacuumresidue, and Arab Medium vacuum residue.

The hydrocracking reactor process produces distillates, hydrocarbongases, hydrogen sulfide and ammonia along with upgraded residues. Thequantities and qualities of these products vary with feedstock, spacevelocity, hydrogen partial pressure and reaction temperature andcatalyst concentration. The upgraded residue yields additionaldistillates when processed in the coker. The combined process ofhydrocracking with the colloidal or molecular catalyst followed bycoking of the hydrocracked residue produces significantly increaseddistillate yields over processing these feeds in either processindependently.

It has been demonstrated that residue conversions as high as 83 wt % canbe achieved by hydrocracking vacuum residue using the colloidal ormolecular catalyst without coking. Without further processing of theresidue, the resultant residue may be sold as low-value bunker fuel.Conversion by coking alone of this same feeds is expected to varybetween 72 and 85 wt %. In each of the 18 examples shown below, theoverall conversion of the combined process is higher than eitherhydrocracking or coking by itself. The overall conversion increase isestimated to be between 4.7 and 15.5 wt %.

Examples 1-5

Examples 1-5 are runs testing the hydroprocessing process as implementedby the invention using as feedstock a blend of 75 wt % Cold Lake and 25wt % Athabasca vacuum residues. Table 1 shows the properties of theindividual vacuum residues, and also of the 75/25 blend that was usedfor the test run.

TABLE 1 Blended Feed for Cold Lake Athabasca Examples 1 to 5 VacuumVacuum (75 wt % Cold Lake/ Resid Resid 25 wt % Athabasca) API Gravity 11.7 C (W %) 82.9 82.12 82.71 H (W %) 9.77 9.64 9.74 S (W %) 5.72 6.515.92 N (W %) 0.8 0.82 0.81 MCR (W %) 22.5 23.1 22.65 Initial BoilingPoint 756.5 657.8 657.8 (° F.) (by TBP Distillation) Resid Content (aswt % 86.44 87.08 86.6 of 975° F.+) (by TBP Distillation)

The results for the pre-coking hydrocracking reactions are based on apilot hydrocracking reactor implemented according to the invention,while the coking reaction results are based on a hypothetical cokingsystem. The pilot hydrocracking reactor system consisted of two reactorstages, each with an internal volume of 2910 cc. Each stage wasback-mixed using an external pumparound loop which continuously withdrewreactor liquid from near the top of the reactor and returned it to thebottom of the reactor. The reactors were operated as bubble columnreactors, without the use of solid supported catalyst. Molybdenum2-ethylhexanoate was used as the catalyst precursor, which was firstblended with a diluent (vacuum gas oil) to generate a diluted catalystprecursor blend. This was then blended into the vacuum residue feedstockmixture described above. Quantities were adjusted so that the dilutedcatalyst precursor blend constituted 1% by weight of the total feedblend, and the catalyst precursor loading in the total feed blend was150 ppm by weight (as Mo). This pre-blended feed mixture was thencharged at regular intervals to a feed vessel, from which it wascontinuously pumped into the two-stage reactor system.

For each of the Examples 1 to 5, performance results are selected from aspecific 24-hour period, which represents the 2^(nd) to 4^(th) day ofoperation on the respective condition, thereby ensuring that the resultsreflect lined-out unit operation. For these examples both reactor stageswere maintained at a common temperature, which differed between examplesas shown in Table 2. Other operating conditions were maintained atconstant values for these examples. Liquid hourly space velocity wasmaintained at 0.37 volume fresh feed/reactor volume/hr, pressure was2200 psig, and the hydrogen feed rate was 4800 scf/bbl fresh feed. Foreach example, the liquid products collected from the unit during thespecified operating day were distilled using a true boiling pointapparatus, to obtain yields and properties for product fractions. Theseresults are shown in Table 2. Of particular note are the properties ofthe vacuum residue product fraction (975° F.+), which provides key datafor the prediction of performance in the hypothetical downstream cokingstep, as discussed below.

TABLE 2 E. 1 Ex. 2 Ex. 3 Ex. 4 Ex. 5 Run Data Operating Period (DayOn-Stream) 4 7 10 15 17 Temperature, F. 814 793 793 784 771 Temperature,C. 435 423 423 418 410 Properties of Resid Product Fraction (975 F+),obtained from TBP Distillation API Gravity −12 −10.2 −10.4 −7.2 −4.7 C(W %) 87 85.95 85.75 85.92 84.85 H (W %) 6.42 7.61 7.77 7.79 8.48 S (W%) 3.52 3.8 3.84 3.9 4.26 N (W %) 2.18 1.84 1.69 1.52 1.54 MCR (W %)62.5 51.4 55.3 49 43.1 Hydrogen Consumption (scf/bbl) 939 1233 1371 1148930 975° F.+ Conversion (ash free) (W %) 72.46 76.51 79.66 72.95 62.51Asphaltene Conversion (HI-TI) (W %) 72.91 78.89 80.15 75.34 69.76 MCRConversion (ash free) (W %) 37.3 54.94 57.01 47.23 40.04 C1-C3 Yield (W%) 4.54 5.04 5.3 4.57 3.37 C4-C7 Yield (W %) 1.79 2.21 2.34 1.79 2.9C4-975 Yield (W %) 68.35 71.76 74.26 69.28 61.6 Resid Yield (975° F.+)(W %) 23.85 20.33 17.61 23.42 32.46

To predict the performance of a downstream coking unit which wouldprocess the vacuum residue product fraction of each of the examples, apublished correlation was used. This was obtained from the book “ProcessChemistry of Petroleum Macromolecues” (I. A. Wiehe, CRC Press, 1^(st)ed., 2008). The theoretical (expected) coke make for a coking operationcan be calculated using the following equation from page 351 of thereference:A(100)=11.28(L)+3.8(100−L)where A is the hydrogen content in the coker feed, 11.28 is the hydrogencontent in the resultant coker derived liquid, 3.8 is the hydrogencontent in the resultant coke, L is the weight percent of the coker feedwhich will become coker derived liquid.

Furthermore, page 390 of the abovementioned reference provides a tableof expected liquid yields for various resid upgrading technologies. Thethree commercial coker technologies listed have liquid yields between 59and 61%, and a fourth demonstrated coking technology is shown with a 66%liquid yield. Based on this data a 65% coker yield was used to calculatethe liquid yield for the present examples.

On this basis, the performance characteristics of the combinedhydrocracking and coking process of the present invention can becompared to those of individual processes. This is shown in Table 3. Theconversion values represent the percentage by weight of hydrocarbonshaving a boiling point of at least 524° C. (975° F.) that are convertedto hydrocarbons having a boiling point below 524° C. (975° F.). For eachexample, the conversion for each individual process (hydrocracking andcoking) is shown alongside that of the combined process of the presentinvention. The column labeled Δ represents the improvement (in absolutepercentage points) of the inventive process relative to the higher ofthe two individual processes. Notably, the combined process improvesconversion rate by about 5 to 12% compared to the best conversion rateprovided by either conventional thermal coking process or hydrocrackingalone.

TABLE 3 Conversion (W %) Coke C4+ Yield (W %) Hydro- Reduct. Hydro-Example cracking Coking Combined Δ (%) cracking Coking Combined Δ 172.46 79.38 84.50 5.13 24.86 68.35 65.00 83.85 15.50 2 76.51 79.38 90.0210.64 51.62 71.76 65.00 84.98 13.22 3 79.66 79.38 91.73 12.08 59.9274.26 65.00 85.71 11.45 4 72.95 79.38 89.07 9.69 47.01 69.28 65.00 84.5015.22 5 62.51 79.38 87.85 8.47 41.07 61.60 65.00 82.70 17.70

Similarly, the yields of C₄+ hydrocarbon distillates having a boilingpoint less than 975° F. are shown in Table 3 for both individualprocesses and the combined process of the present invention. ForExamples 1 to 5, the combined process improves the C₄+ distillate yieldby 11.45-17.7 absolute percentage points, as compared to the best C₄+distillate yield obtained by either process individually. Complementaryto the conversion rate and C₄+ yield results, the combined hydrocrackingand thermal cracking process substantially reduces coke formation,lowering coke formation by 24.86-59.92%.

Example 6-9

Examples 6-9 were conducted in similar fashion to the abovementionedexamples, with hydrocracking tests conducted in a continuous-flow pilotunit, and coking performance predicted using the methods of the citedliterature reference. For these examples, several different feedstockcompositions were tested. For Example 6, Black Rock atmospheric residuewas tested. Examples 7 and 8 utilized Black Rock vacuum residue. Example9 used Murphy vacuum residue. The properties of each of these feedmaterials are summarized in Table 4.

TABLE 4 Black Rock Black Rock Murphy Atmospheric Vacuum Vacuum ResidueResidue Residue Examples 6 7 and 8 9 API Gravity 6.7 1.4 0.6 C (W %)81.27 80.32 80.44 H (W %) 10.12 9.26 9.35 S (W %) 7.29 8.8 8.6 N (W %)0.64 0.87 0.81 MCR (W %) 13.87 26.55 24.21 Initial Boiling Point (° F.)412 n/a n/a (by TBP Distillation) Resid Content (as wt % of 59.24 95.2691.05 975° F.+) (by TBP Distillation)

Test conditions for Examples 6-9 differed in several respects from thosein the earlier examples. The hydrocracking test unit consisted of only asingle reactor stage of 2910 cc internal volume. In addition to changesin temperature between the individual examples, there were alsovariations in pressure, liquid hourly space velocity, hydrogen feedrate, and catalyst concentration; these are shown by condition in Table5. Table 5 also shows the hydrocracking performance results for each ofthe examples.

TABLE 5 Ex. 6 Ex. 7 Ex. 8 Ex. 9 Run Data Operating Period (DayOn-Stream) 9 18 24 28 Temperature (° F.) 806 833 819 818 Temperature (°C.) 430 445 437 437 Pressure (psig) 2000 2400 2400 2400 LHSV (vol freshfeed/vol 0.8 0.39 0.39 0.39 reactor/hr) Hydrogen Feed (scf/bbl freshfeed) 3999 5697 5098 5142 Catalyst Concentration 65 250 175 225 (ppmw asMo) Properties of Resid Product Fraction (975 F+), obtained from TBPDistillation API Gravity −11.6 −9.3 −6.8 −9 C (W %) 83.91 84.54 84.3784.68 H (W %) 7.67 7.54 7.91 7.69 S (W %) 5.73 4.67 5.14 4.51 N (W %)1.32 1.62 1.38 1.47 MCR (W %) 39.09 54.09 48.52 45.24 HydrocrackingProcess Performance Hydrogen Consumption (scf/bbl) 667 1973 1453 1718975° F.+ Conversion 61.53 82.77 73.28 79.75 (ash free) (W %) AsphalteneConversion (HI-TI) 58.47 78.3 69.36 78.56 (W %) MCR Conversion (ashfree) (W %) 36.89 62.88 52.58 65.04 Yields, moisture and ash-free freshfeed basis C1-C3 Yield (W %) 3.04 7.2 5.7 7.02 C4-C7 Yield (W %) 1.6 3.52.96 2.93 C4-975 Yield (W %) 71 71.43 64.22 69.38 Resid Yield (975° F.+)(W %) 22.28 16.01 24.82 18.36

Similar to the previous examples, the performance of a downstream cokingunit was predicted for Examples 6 to 9 based on the results ofhydrocracking tests in Table 5 and the coking correlation obtained fromthe literature. Table 6 shows the results, comparing the 524° C.+(975°F.+) vacuum residue conversion for the individual processes(hydrocracking and coking) to that of the combined process of thepresent invention. The combined process improves conversion rate by 4.76to 15.54 absolute percentage points compared to the best conversion rateobtained by either coking or hydrocracking individually.

TABLE 6 Conversion (W %) Coke C4+ Yield (W %) Hydro- Reduct. Hydro-Example cracking Coking Combined Δ (%) cracking Coking Combined Δ 661.53 84.49 89.25 4.76 30.67 72.60 65.00 87.08 14.48 7 82.77 72.99 92.009.23 70.36 74.92 65.00 85.33 10.41 8 73.28 72.99 88.82 15.54 58.59 67.1865.00 83.31 16.14 9 79.75 74.20 91.19 11.44 65.84 72.32 65.00 84.2511.94

Similarly, the C4+ distillate yield is compared between the individualprocesses and the combined process of the present invention. Thecombined process improves C4+ distillate yield by 10.41-16.14 percentagepoints, as compared to the best C₄+ distillate yield by either processindividually. Complementary to the conversion rate and C4+ yieldresults, the process combining hydrocracking and thermal crackingsubstantially reduces coke formation, lowering coke formation by30.67-70.36%.

Example 10-13

Examples 10 to 13 were conducted in a fashion similar to the previousexamples. In this case, all examples used a Ku 850° F.+ vacuum residue,the properties of which are shown in Table 7.

TABLE 7 Ku 850° F.+ Vacuum Residue Examples 10 to 13 API Gravity 3 C (W%) 82.31 H (W %) 9.64 S (W %) 6.11 N (W %) 1.02 MCR (W %) 25.41 InitialBoiling Point (° F.) 694 (by D-1160 Distillation) Resid Content (as wt %of 975° F.+) 82.23 (by D-1160 Distillation)

For Examples 10-13, the hydrocracking pilot unit consisted of a singlereactor stage of 2910 cc internal volume. Most operating conditions weremaintained at constant levels for these examples, with reactortemperature at 815° F. (435° C.), pressure at 2500 psig, hydrogen feedrate at 5600 scf/bbl fresh feed, and liquid hourly space velocity at0.35 volume fresh feed/volume reactor/hr. The only difference betweenthese examples is the catalyst concentration, which ranged from 51 to508 ppm by weight (as Mo), as shown in Table 8.

TABLE 8 Ex. 10 Ex. 11 Ex. 12 Ex. 13 Run Data Operating Period (DayOn-Stream) 4 9 13 16 Catalyst Concentration 153 153 508 51 (ppmw as Mo)Properties of Resid Product Fraction (975 F+), obtained from TBPDistillation API Gravity −8.1 −8.1 −8.5 −9.1 C (W %) 84.68 84.94 84.6584.84 H (W %) 7.81 7.68 7.85 7.66 S (W %) 4.55 4.81 4.25 4.46 N (W %)1.36 1.44 1.46 1.32 MCR (W %) 50.95 53.79 53.58 51.98 HydrocrackingProcess Performance Hydrogen Consumption (scf/bbl) 1685 1656 1605 1632975° F.+ Conversion (ash free) 75.3 74.55 75.16 74.86 (W %) AsphalteneConversion (HI-TI) 76.03 74.89 78.85 72.05 (W %) MCR Conversion (ashfree) (W %) 53.89 53.65 53.72 53.67 Yields, moisture and ash-free freshfeed basis C1-C3 (W %) 7.33 7.26 6.86 7.21 C4-350° F. (W %) 16.48 16.6716.72 17.28 350-650° F. (W %) 26.12 26.44 26.7 26.57 650-975° F. (W %)26.65 25.58 26.06 24.94 Resid Yield 975° F.+ (W %) 20.3 20.92 20.4220.66

Table 8 also shows the performance results for the hydrocrackingprocess, including the properties of the vacuum residue product fraction(975F+), the process residue conversion, and the yields of productfractions. As was done for the previous examples, the performance of adownstream coking process was then predicted using the correlationobtained from the abovementioned literature reference, allowing theperformance of the individual processes (hydrocracking and coking) to becompared to that of the combined process of the present invention.Results are shown in Table 9. For the vacuum residue conversion, thecombined process increases conversion by 11.85 to 15.54 percentagepoints compared to the best result obtained from either individualprocess.

TABLE 9 Conversion (W %) Coke C4+ Yield (W %) Hydro- Reduct. Hydro-Example cracking Coking Combined Δ (%) cracking Coking Combined Δ 1075.30 78.07 90.58 12.51 57.04 69.25 65.00 82.44 13.20 11 74.55 78.0789.93 11.85 54.07 68.70 65.00 82.30 13.60 12 75.16 78.07 90.64 12.5657.3 69.48 65.00 82.75 13.27 13 74.86 78.07 90.00 11.93 54.39 68.7965.00 82.23 13.43

Similarly, the C4+ distillate yield is shown for the individualprocesses and the combined process of the present invention. Thecombined process improves C4+ distillate yield by 13.27-13.60 percentagepoints, as compared to the best result obtained from either processindividually. Complementary to the conversion rate and C4+ yieldresults, the process combining hydrocracking and thermal crackingsubstantially reduces coke formation, lowering coke formation by54.07-57.3%.

Example 14-18

Examples 14 to 18 were conducted in similar fashion to the previousexamples. In this case, the feedstock was Arab Medium vacuum residue,the properties of which are shown in Table 10.

TABLE 10 Arab Medium Vacuum Residue Examples 14 to 18 API Gravity 6.48 C(W %) 83.1 H (W %) 10.04 S (W %) 5.07 N (W %) 0.67 MCR (W %) 17.21Initial Boiling Point (° F.) 522 (by D-1160 Distillation) Resid Content(as wt % of 1000° F.+) 92.65 (by D-1160 Distillation)

For Examples 14 to 18, the hydrocracking pilot unit consisted of a tworeactor stages of 2910 cc internal volume each. For these examples,reactor temperature was maintained at 803° F. (428° C.) and pressure at2250 psig. Other conditions were varied as shown in Table 11.

TABLE 11 Ex. 14 Ex. 15 Ex. 16 Ex. 17 Ex. 18 Run Data Operating Period(Day On-Stream) 4 10 14 27 32 LHSV (vol fresh feed/vol reactor/hr) 0.390.29 0.22 0.29 0.29 Hydrogen Feed (scf/bbl fresh feed) 4511 4513 47044515 4505 Catalyst Concentration (ppmw as Mo) 149 149 149 298 50Properties of Resid Product Fraction (975 F+), obtained from TBPDistillation API Gravity −2.2 −6.3 −7.7 −3.9 −5.6 C (W %) 84.52 85.385.04 85.98 85 H (W %) 8.72 8.3 7.83 8.54 8.24 S (W %) 4.32 4.25 3.993.85 4.49 N (W %) 1.41 1.58 1.63 1.3 1.62 MCR (W %) 41.53 47.46 42.0445.33 42.46 Hydrocracking Process Performance Hydrogen Consumption(scf/bbl) 838 1021 1386 1249 1149 1000° F.+ Conversion (ash free) (W %)63.13 73.4 80.74 73.44 73.68 Asphaltene Conversion (HI-TI) (W %) 37.849.14 56.66 55.65 42.84 MCR Conversion (ash free) (W %) 15.84 29.28 38.129.07 30.76 Yields, moisture and ash-free fresh feed basis C1-C3 (W %)ff 3.47 5 6.35 5.04 5.17 C4-650° F. (W %) ff 31.29 39.68 48.31 40.0141.93 650-1000° F. (W %) ff 28.72 27.78 24.76 27.7 25.97 Resid Yield1000° F.+ (W %) ff 34.33 24.77 17.93 24.73 24.51

Table 11 also shows the performance results for the hydrocrackingprocess, including the properties of the vacuum residue product fraction(1000° F.+), the process residue conversion, and the yields of productfractions. As was done for the previous examples, the performance of adownstream coking process was then predicted using the correlationobtained from the abovementioned literature reference, allowing theperformance of the individual processes (hydrocracking and coking) to becompared to that of the combined process of the present invention.Results are shown in Table 12. For the vacuum residue conversion, thecombined process improves conversion rate by 4.83 to 8.31 percentagepoints compared to the best result obtained from either processindividually.

TABLE 12 Conversion (W %) Coke C4+ Yield (W %) Hydro- Reduct. Hydro-Example cracking Coking Combined Δ (%) cracking Coking Combined Δ 1463.13 83.42 88.25 4.83 29.12 60.01 65.00 82.33 17.33 15 73.40 83.4290.13 6.71 40.47 67.46 65.00 83.56 16.10 16 80.74 83.42 91.73 8.31 50.173.07 65.00 84.73 11.66 17 73.44 83.42 90.94 7.52 45.35 67.71 65.0083.78 16.08 18 73.68 83.42 90.04 6.62 39.91 67.90 65.00 83.83 15.93

Similarly, the C4+ distillate yield is compared between the individualprocesses and the combined process of the present invention. Thecombined process improves C4+ distillate yield by 11.66-17.33 percentagepoints, as compared to the best yield provided by either processindividually. Complementary to the conversion rate and C4+ yieldresults, the process combining hydrocracking and thermal crackingsubstantially reduces coke formation, lowering coke formation by29.12-45.35%.

The present invention may be embodied in other specific forms withoutdeparting from its spirit or essential characteristics. The describedembodiments are to be considered in all respects only as illustrativeand not restrictive. The scope of the invention is, therefore, indicatedby the appended claims rather than by the foregoing description. Allchanges which come within the meaning and range of equivalency of theclaims are to be embraced within their scope.

What is claimed is:
 1. A method for hydroprocessing a heavy oilfeedstock to increase production of upgraded liquid hydrocarbon productsand reduce coke formation, the method comprising: preparing a heavy oilfeedstock comprised of a substantial quantity of hydrocarbons having aboiling point greater than about 343° C. (650° F.), includingasphaltenes or other coke forming precursors, and a colloidal ormolecular catalyst formed in situ within and dispersed throughout theheavy oil feedstock; introducing the heavy oil feedstock and hydrogeninto a pre-coking hydrocracking reactor; hydrocracking the heavy oilfeedstock at hydrocracking conditions to cause fragmentation of largerhydrocarbon molecules into smaller molecular fragments having a fewernumber of carbon atoms and form hydrocarbon free radicals from the heavyoil feedstock, the colloidal or molecular catalyst catalyzing upgradingreactions between hydrogen and the hydrocarbon free radicals to yield anupgraded material, the upgrading reactions reducing the quantity ofasphaltenes or other coke forming precursors by at least 20%, increasingthe hydrogen to carbon ratio in the upgraded material, and decreasingthe boiling points of hydrocarbons in the upgraded material compared tothe heavy oil feedstock; transferring the upgraded material, togetherwith residual colloidal or molecular catalyst and hydrogen, to aseparator to separate gaseous and volatile fractions from a liquidhydrocarbon fraction, the liquid hydrocarbon fraction comprisingresidual catalyst metal; introducing at least a portion of the liquidhydrocarbon fraction and residual catalyst metal into one or more cokingreactors and causing thermal-cracking of the liquid hydrocarbon fractionto form upgraded hydrocarbon products and coke; separating the coke fromthe upgraded hydrocarbon products; and further processing the upgradedhydrocarbon products to form further processed hydrocarbons and withoutrecycling any further processed hydrocarbons to the hydrocrackingreactor.
 2. The method as in claim 1, wherein the separator comprises ahot separator.
 3. The method as in claim 1, wherein the separatorcomprises a distillation tower.
 4. The method as in claim 1, wherein theseparator comprises a hot separator and a distillation tower.
 5. Themethod as in claim 1, wherein the portion of the liquid hydrocarbonfraction introduced into the coking reactor comprises a vacuum reducedcrude.
 6. The method as defined in claim 1, the pre-coking hydrocrackingreactor comprising at least one of a slurry phase reactor, an ebullatedbed reactor, or a fixed bed reactor.
 7. The method as in 1, wherein thepre-coking hydrocracking reactor is a slurry phase reactor including (i)an inlet port at a bottom of the slurry phase reactor into which theheavy oil feedstock and hydrogen are introduced and (ii) an outlet portat a top of the slurry phase reactor from which the upgraded material,colloidal or molecular catalyst, and hydrogen are withdrawn.
 8. Themethod as defined in claim 7, the slurry phase reactor furthercomprising a recycle channel, a recycling pump, and a distributor gridplate.
 9. The method as in claim 1, wherein the one or more cokingreactors comprise one or more delayed coking reactors, fluid cokingreactors, or Flexicoking® reactors.
 10. The method as defined in claim1, wherein the heavy oil feedstock comprises at least about 10 wt %asphaltenes or other coke forming precursors.
 11. The method as in claim1, the upgrading reactions reducing the quantity of asphaltenes or othercoke forming precursors by at least 40 wt %.
 12. The method as in claim1, the upgrading reactions reducing the quantity of asphaltenes or othercoke forming precursors by at least 60 wt %.
 13. The method as in claim1, the method converting at least 60 wt % of hydrocarbons having aboiling point of at least 524° C. (975° F.).
 14. The method as in claim1, the method converting at least 70 wt % of hydrocarbons having aboiling point of at least 524° C. (975° F.).
 15. The method as in claim1, the method converting at least 80 wt % of hydrocarbons having aboiling point of at least 524° C. (975° F.).
 16. The method as in claim1, wherein the method converts at least 85 wt % of hydrocarbons having aboiling point of at least 524° C. (975° F.).
 17. The method as in claim1, wherein the method yields at least 80 wt % of C4+hydrocarbons and aboiling point of less than about 524° C. (975° F.).
 18. The method as inclaim 1, wherein the method reduces coke formation by at least 25 wt %compared to coking in the absence of hydrocracking catalyzed by themolecular or colloidal catalyst.
 19. The method as defined in claim 1,the colloidal or molecular catalyst in the heavy oil feedstock beingformed by: mixing a hydrocarbon oil diluent and an oil soluble catalystprecursor at a temperature below which a significant portion of thecatalyst precursor starts to decompose to form a diluted precursormixture; mixing the diluted precursor mixture with a heavy oil feedstockin a manner so as to yield a conditioned feedstock that forms thecolloidal or molecular catalyst upon decomposing the catalyst precursorand allowing metal liberated therefrom to react with sulfur liberatedfrom the feedstock; and heating the conditioned feedstock so as todecompose the catalyst precursor and allow metal liberated from thedecomposed catalyst precursor to react with sulfur liberated from theheavy oil feedstock so as to form the colloidal or molecular catalyst.20. The method as defined in claim 19, the hydrocarbon oil diluentcomprising at least one of vacuum gas oil, decant oil, cycle oil, orlight gas oil.
 21. The method as defined in claim 19, the catalystprecursor comprising at least one transition metal and at least oneorganic moiety comprising or derived from 3-cyclopentylpropionic acid,cyclohexanebutyric acid, biphenyl-2-carboxylic acid, 4-heptylbenzoicacid, 5-phenylvaleric acid, geranic acid, 10-undecenoic acid, dodecanoicacid, octanoic acid, 2-ethylhexanoic acid, naphthanic acid,pentacarbonyl, or hexacarbonyl.
 22. The method as defined in claim 21,the at least one transition metal comprising one or more of Mo, Ni, Co,W, V or Fe.
 23. The method as defined in claim 19, the catalystprecursor comprising at least one of molybdenum 3-cyclopentylpropionate,molybdenum cyclohexanebutanoate, molybdenum biphenyl-2-carboxylate,molybdenum 4-heptylbenzoate, molybdenum 5-phenylpentanoate, molybdenumgeranate, molybdenum 10-undecenoate, molybdenum dodecanoate, molybdenum2-ethylhexanoate, molybdenum naphthanate, molybdenum hexacarbonyl,vanadium octoate, vanadium naphthanate, or iron pentacarbonyl.
 24. Themethod as defined in claim 19, the ratio of catalyst precursorcomposition to hydrocarbon oil diluent being in a range of about 1:100to about 1:5.
 25. The method as defined in claim 19, the hydrocarbon oildiluent and catalyst precursor composition being mixed at temperature ina range of about 25° C. to about 250° C., the diluted precursor mixtureand heavy oil feedstock being mixed at a temperature in a range of about25° C. to about 350° C., and the conditioned feedstock being heated to atemperature in a range of about 275° C. to about 375° C.
 26. The methodas defined in claim 19, the hydrocarbon oil diluent and catalystprecursor composition being mixed at temperature in a range of about 75°C. to about 150° C., the diluted precursor mixture and heavy oilfeedstock being mixed at a temperature in a range of about 75° C. toabout 250° C., and the conditioned feedstock being heated to atemperature in a range of about 310° C. to about 360° C.
 27. The methodas defined in claim 19, the hydrocarbon oil diluent and catalystprecursor composition being mixed for a time period in a range of about1 second to about 20 minutes, and the diluted precursor mixture andheavy oil feedstock being mixed for a time period in a range of about 1second to about 20 minutes.
 28. The method as defined in claim 19, thehydrocarbon oil diluent and catalyst precursor composition being mixedfor a time period in a range of about 20 seconds to about 3 minutes, andthe diluted precursor mixture and heavy oil feedstock being mixed for atime period in a range of about 20 seconds to about 5 minutes.
 29. Themethod as defined in claim 1, further comprising processing the upgradedhydrocarbon products downstream from the pre-coking hydrocrackingreactor and the one or more coking reactors.
 30. A method forhydroprocessing a heavy oil feedstock to increase production of upgradedliquid hydrocarbon products and reduce coke formation, the methodcomprising: preparing a heavy oil feedstock comprised of a substantialquantity of hydrocarbons having a boiling point greater than about 343°C. (650° F.) and including asphaltenes, and a colloidal or molecularcatalyst formed in situ within and dispersed throughout the heavy oilfeedstock, wherein the heavy oil feedstock is selected from the groupconsisting of heavy crude oil, oil sand bitumen, atmospheric towerbottoms, vacuum tower bottoms, resid, visbreaker bottoms, coal tar,heavy oil from oil shale, and liquefied coal, with the proviso that theheavy oil feedstock does not comprise coker gas oil; introducing theheavy oil feedstock and hydrogen into a pre-coking hydrocrackingreactor; hydrocracking the heavy oil feedstock at hydrocrackingconditions to cause fragmentation of at least some of the asphaltenesand to form hydrocarbon free radicals from the heavy oil feedstock, thecolloidal or molecular catalyst catalyzing upgrading reactions betweenhydrogen and the hydrocarbon free radicals to yield an upgradedmaterial, the upgrading reactions reducing the quantity of asphaltenesby at least 20%, increasing the hydrogen to carbon ratio in the upgradedmaterial, and decreasing the boiling points of hydrocarbons in theupgraded material compared to the heavy oil feedstock; transferring theupgraded material, together with residual colloidal or molecularcatalyst and hydrogen, to a separator to separate gaseous and volatilefractions from a liquid hydrocarbon fraction, the liquid hydrocarbonfraction comprising residual catalyst metal; introducing at least aportion of the liquid hydrocarbon fraction and residual catalyst metalinto one or more coking reactors and causing thermal-cracking of theliquid hydrocarbon fraction to form upgraded hydrocarbon products andcoke; separating the coke from the upgraded hydrocarbon products, theupgraded hydrocarbon products including coker gas oil; and furtherprocessing an entirety of the upgraded hydrocarbon products, includingthe coker gas oil, downstream from the pre-coking hydrocracking reactorto form one or more further processed hydrocarbons.
 31. The method asdefined in claim 30, wherein the heavy oil feedstock consistsessentially of heavy crude oil, oil sand bitumen, atmospheric towerbottoms, vacuum tower bottoms, resid, visbreaker bottoms, coal tar,heavy oil from oil shale, or liquefied coal.
 32. The method as in claim30, the upgrading reactions reducing the quantity of asphaltenes by atleast 40 wt %.
 33. The method as in claim 30, the upgrading reactionsreducing the quantity of asphaltenes by at least 60 wt %.
 34. The methodas in claim 1, wherein the heavy oil feedstock comprises at least one ofheavy crude oil, oil sand bitumen, atmospheric tower bottoms, vacuumtower bottoms, resid, visbreaker bottoms, coal tar, heavy oil from oilshale, or liquefied coal.
 35. The method as in claim 1, wherein thehydrocracking temperature is in a range of about 395° C. to about 460°C., the upgrading reactions reducing the quantity of asphaltenes orother coke forming precursors by at least 20 wt %.
 36. The method as inclaim 30, wherein the hydrocracking temperature is in a range of about395° C. to about 460° C., the upgrading reactions reducing the quantityof asphaltenes or other coke forming precursors by at least 20 wt %.